Cs is the concentration of impurity or minor component in the solid
phase, and Cᐉ is the impurity concentration in the liquid phase. The
distribution coefficient generally varies with composition. The value
of k is greater than 1 when the solute raises the melting point and less
than 1 when the melting point is depressed. In the regions near pure
A or B the liquidus and solidus lines become linear; i.e., the distribu-
tion coefficient becomes constant. This is the basis for the common
assumption of constant k in many mathematical treatments of frac-
tional solidification in which ultrapure materials are obtained.
In the case of a simple eutectic system shown in Fig. 20-2, a pure
solid phase is obtained by cooling if the composition of the feed mix-
ture is not at the eutectic composition. If liquid composition is eutec-
tic, then separate crystals of both species will form. In practice it is
difficult to attain perfect separation of one component by crystalliza-
tion of a eutectic mixture. The solid phase will always contain trace
amounts of impurity because of incomplete solid-liquid separation,
slight solubility of the impurity in the solid phase, or volumetric inclu-
sions. It is difficult to generalize on which of these mechanisms is the
major cause of contamination because of analytical difficulties in the
The distribution-coefficient concept is commonly applied to frac-
tional solidification of eutectic systems in the ultrapure portion of the
phase diagram. If the quantity of impurity entrapped in the solid
phase for whatever reason is proportional to that contained in the
melt, then assumption of a constant k is valid. It should be noted that
the theoretical yield of a component exhibiting binary eutectic behav-
ior is fixed by the feed composition and position of the eutectic. Also,
in contrast to the case of a solid solution, only one component can be
obtained in a pure form.
There are many types of phase diagrams in addition to the two cases
presented here; these are summarized in detail by Zief and Wilcox
(op. cit., p. 21). Solid-liquid phase equilibria must be determined
experimentally for most binary and multicomponent systems. Predic-
tive methods are based mostly on ideal phase behavior and have
limited accuracy near eutectics. A predictive technique based on
extracting liquid-phase activity coefficients from vapor-liquid equilib-
ria that is useful for estimating nonideal binary or multicomponent
solid-liquid phase behavior has been reported by Muir (Pap. 71f, 73d
ann. meet., AIChE, Chicago, 1980).
Progressive freezing, sometimes called normal freezing, is the slow,
directional solidification of a melt. Basically, this involves slow solidifi-
cation at the bottom or sides of a vessel or tube by indirect cooling.
The impurity is rejected into the liquid phase by the advancing solid
interface. This technique can be employed to concentrate an impurity
or, by repeated solidifications and liquid rejections, to produce a very
pure ingot. Figure 20-3 illustrates a progressive freezing apparatus.
The solidification rate and interface position are controlled by the rate
of movement of the tube and the temperature of the cooling medium.
There are many variations of the apparatus; e.g., the residual-liquid
portion can be agitated and the directional freezing can be carried out
vertically as shown in Fig. 20-3 or horizontally (see Richman et al., in
Zief and Wilcox, op. cit., p. 259). In general, there is a solute redistri-
bution when a mixture of two or more components is directionally
Component Separation by Progressive Freezing When the
distribution coefficient is less than 1, the first solid which crystallizes
contains less solute than the liquid from which it was formed. As the
fraction which is frozen increases, the concentration of the impurity in
the remaining liquid is increased and hence the concentration of
impurity in the solid phase increases (for k < 1). The concentration
gradient is reversed for k > 1. Consequently, in the absence of diffu-
sion in the solid phase a concentration gradient is established in the
One extreme of progressive freezing is equilibrium freezing. In this
case the freezing rate must be slow enough to permit diffusion in the
solid phase to eliminate the concentration gradient. When this occurs,
there is no separation if the entire tube is solidified. Separation can be
achieved, however, by terminating the freezing before all the liquid
has been solidified. Equilibrium freezing is rarely achieved in practice
because the diffusion rates in the solid phase are usually negligible
(Pfann, op. cit., p. 10).
If the bulk-liquid phase is well mixed and no diffusion occurs in the
solid phase, a simple expression relating the solid-phase composition
to the fraction frozen can be obtained for the case in which the distri-
bution coefficient is independent of composition and fraction frozen
[Pfann, Trans. Am. Inst. Mech. Eng., 194, 747 (1952)].
Cs = kC0(1 − X)k − 1
C0 is the solution concentration of the initial charge, and X is the frac-
tion frozen. Figure 20-4 illustrates the solute redistribution predicted
by Eq. (20-2) for various values of the distribution coefficient.
There have been many modifications of this idealized model to
account for variables such as the freezing rate and the degree of
mixing in the liquid phase. For example, Burton et al. [J. Chem. Phys.,
21, 1987 (1953)] reasoned that the solid rejects solute faster than it
can diffuse into the bulk liquid. They proposed that the effect of the
20-4 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-2 Simple eutectic-phase diagram at constant pressure. (Zief and
Wilcox, Fractional Solidification, vol. 1, Marcel Dekker, New York, 1967, p. 24.)
FIG. 20-3 Progressive freezing apparatus.
freezing rate and stirring could be explained by the diffusion of solute
through a stagnant film next to the solid interface. Their theory
resulted in an expression for an effective distribution coefficient keff
which could be used in Eq. (20-2) instead of k.
keff = (20-3)
where f = crystal growth rate, cm/s; δ = stagnant film thickness, cm;
and D = diffusivity, cm2
/s. No further attempt is made here to sum-
marize the various refinements of Eq. (20-2). Zief and Wilcox (op. cit.,
p. 69) have summarized several of these models.
Pertinent Variables in Progressive Freezing The dominant
variables which affect solute redistribution are the degree of mixing in
the liquid phase and the rate of solidification. It is important to attain
sufficient mixing to facilitate diffusion of the solute away from the solid-
liquid interface to the bulk liquid. The film thickness δ decreases as the
level of agitation increases. Cases have been reported in which essen-
tially no separation occurred when the liquid was not stirred. The freez-
ing rate which is controlled largely by the lowering rate of the tube (see
Fig. 20-3) has a pronounced effect on the separation achieved. The sep-
aration is diminished as the freezing rate is increased. Also fluctuations
in the freezing rate caused by mechanical vibrations and variations in
the temperature of the cooling medium can decrease the separation.
Applications Progressive freezing has been applied to both solid
solution and eutectic systems. As Fig. 20-4 illustrates, large separation
factors can be attained when the distribution coefficient is favorable.
Relatively pure materials can be obtained by removing the desired
portion of the ingot. Also in some cases progressive freezing provides
a convenient method of concentrating the impurities; e.g., in the case
of k < 1 the last portion of the liquid that is frozen is enriched in the
Progressive freezing has been applied on the commercial scale. For
example, aluminum has been purified by continuous progressive
freezing [Dewey, J. Metals, 17, 940 (1965)]. The Proabd refiner de-
scribed by Molinari (Zief and Wilcox, op. cit., p. 393) is also a com-
mercial example of progressive freezing. In this apparatus the mixture
is directionally solidified on cooling tubes. Purification is achieved
1 + (1/k − 1)e−f δ/ D
because the impure fraction melts first; this process is called sweating.
This technique has been applied to the purification of naphthalene
and p-dichlorobenzene and commercial equipment is available from
BEFS PROKEM, Houston, Tx.
Zone melting also relies on the distribution of solute between the liq-
uid and solid phases to effect a separation. In this case, however, one
or more liquid zones are passed through the ingot. This extremely ver-
satile technique, which was invented by W. G. Pfann, has been used to
purify hundreds of materials. Zone melting in its simplest form is illus-
trated in Fig. 20-5. A molten zone can be passed through an ingot
from one end to the other by either a moving heater or by slowly draw-
ing the material to be purified through a stationary heating zone.
Progressive freezing can be viewed as a special case of zone melt-
ing. If the zone length were equal to the ingot length and if only one
pass were used, the operation would become progressive freezing. In
general, however, when the zone length is only a fraction of the ingot
length, zone melting possesses the advantage that a portion of the
ingot does not have to be discarded after each solidification. The last
portion of the ingot which is frozen in progressive freezing must be
discarded before a second freezing.
Component Separation by Zone Melting The degree of
solute redistribution achieved by zone melting is determined by the
zone length l, ingot length L, number of passes n, the degree of mix-
ing in the liquid zone, and the distribution coefficient of the materials
being purified. The distribution of solute after one pass can be
obtained by material-balance considerations. This is a two-domain
problem; i.e., in the major portion of the ingot of length L − l zone
melting occurs in the conventional sense. The trailing end of the ingot
of length l undergoes progressive freezing. For the case of constant-
distribution coefficient, perfect mixing in the liquid phase, and negli-
gible diffusion in the solid phase, the solute distribution for a single
pass is given by Eq. (20-4) [Pfann, Trans. Am. Inst. Mech. Eng., 194,
Cs = C0[1 − (1 − k)e−kx/ᐉ
The position of the zone x is measured from the leading edge of the
ingot. The distribution for multiple passes can also be calculated from
a material balance, but in this case the leading edge of the zone
encounters solid corresponding to the composition at the point in
question for the previous pass. The multiple-pass distribution has
been numerically calculated (Pfann, Zone Melting, 2d ed., Wiley, New
York, 1966, p. 285) for many combinations of k, L/l, and n. Typical
solute-composition profiles are shown in Fig. 20-6 for various num-
bers of passes.
The ultimate distribution after an infinite number of passes is also
shown in Fig. 20-6 and can be calculated for x < (L − l) from the fol-
lowing equation (Pfann, op. cit., p. 42):
Cs = AeBX
where A and B can be determined from the following relations:
k = Bᐉ/(eBᐉ
− 1) (20-6)
A = C0BL/(eBL
− 1) (20-7)
CRYSTALLIZATION FROM THE MELT 20-5
FIG. 20-4 Curves for progressive freezing, showing solute concentration C in
the solid versus fraction-solidified X. (Pfann, Zone Melting, 2d ed., Wiley, New
York, 1966, p. 12.)
FIG. 20-5 Diagram of zone refining.
The ultimate distribution represents the maximum separation that can
be attained without cropping the ingot. Equation (20-5) is approxi-
mate because it does not include the effect of progressive freezing in
the last zone length.
As in progressive freezing, many refinements of these models have
been developed. Corrections for partial liquid mixing and a variable
distribution coefficient have been summarized in detail (Zief and
Wilcox, op. cit., p. 47).
Pertinent Variables in Zone Melting The dominant variables
in zone melting are the number of passes, ingot-length–zone-length
ratio, freezing rate, and degree of mixing in the liquid phase. Figure
20-6 illustrates the increased solute redistribution that occurs as the
number of passes increases. Ingot-length–zone-length ratios of 4 to 10
are commonly used (Zief and Wilcox, op. cit., p. 624). An exception is
encountered when one pass is used. In this case the zone length
should be equal to the ingot length; i.e., progressive freezing provides
the maximum separation when only one pass is used.
The freezing rate and degree of mixing have effects in solute redis-
tribution similar to those discussed for progressive freezing. Zone
travel rates of 1 cm/h for organic systems, 2.5 cm/h for metals, and
20 cm/h for semiconductors are common. In addition to the zone-
travel rate the heating conditions affect the freezing rate. A detailed
summary of heating and cooling methods for zone melting has been
outlined by Zief and Wilcox (op. cit., p. 192). Direct mixing of the liq-
uid region is more difficult for zone melting than progressive freezing.
Mechanical stirring complicates the apparatus and increases the prob-
ability of contamination from an outside source. Some mixing occurs
because of natural convection. Methods have been developed to stir
the zone magnetically by utilizing the interaction of a current and
magnetic field (Pfann, op. cit., p. 104) for cases in which the charge
material is a reasonably good conductor.
Applications Zone melting has been used to purify hundreds of
inorganic and organic materials. Many classes of inorganic compounds
including semiconductors, intermetallic compounds, ionic salts, and
oxides have been purified by zone melting. Organic materials of many
types have been zone-melted. Zief and Wilcox (op. cit., p. 624) have
compiled tables which give operating conditions and references for
both inorganic and organic materials with melting points ranging from
−115°C to over 3000°C.
Some materials are so reactive that they cannot be zone-melted to a
high degree of purity in a container. Floating-zone techniques in
which the molten zone is held in place by its own surface tension have
been developed by Keck et al. [Phys. Rev., 89, 1297 (1953)].
Continuous-zone-melting apparatus has been described by Pfann
(op. cit., p. 171). This technique offers the advantage of a close
approach to the ultimate distribution, which is usually impractical for
Performance data have been reported by Kennedy et al. (The
Purification of Inorganic and Organic Materials, Marcel Dekker, New
York, 1969, p. 261) for continuous-zone refining of benzoic acid.
MELT CRYSTALLIZATION FROM THE BULK
Conducting crystallization inside a vertical or horizontal column with
a countercurrent flow of crystals and liquid can produce a higher
product purity than conventional crystallization or distillation. The
working concept is to form a crystal phase from the bulk liquid, either
internally or externally, and then transport the solids through a coun-
tercurrent stream of enriched reflux liquid obtained from melted
product. The problem in practicing this technology is the difficulty of
controlling solid-phase movement. Unlike distillation, which exploits
the specific-gravity differences between liquid and vapor phases, melt
crystallization involves the contacting of liquid and solid phases that
have nearly identical physical properties. Phase densities are fre-
quently very close, and gravitational settling of the solid phase may be
slow and ineffective. The challenge of designing equipment to accom-
plish crystallization in a column has resulted in a myriad of configura-
tions to achieve reliable solid-phase movement, high product yield
and purity, and efficient heat addition and removal.
Investigations Crystallization conducted inside a column is cate-
gorized as either end-fed or center-fed depending on whether the
feed location is upstream or downstream of the crystal forming section.
Figure 20-7 depicts the features of an end-fed commercial column
described by McKay et al. [Chem. Eng. Prog. Symp. Ser., no. 25, 55, 163
(1969)] for the separation of xylenes. Crystals of p-xylene are formed by
indirect cooling of the melt in scraped-surface heat exchangers, and the
resultant slurry is introduced into the column at the top. This type of col-
umn has no mechanical internals to transport solids and instead relies
upon an imposed hydraulic gradient to force the solids through the col-
umn into the melting zone. Residue liquid is removed through a filter
directly above the melter. A pulse piston in the product discharge
improves washing efficiency and column reliability.
20-6 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-6 Relative solute concentration C/C0 (logarithmic scale) versus dis-
tance in zone lengths x/ᐉ from beginning of charge, for various numbers of
passes n. L denotes charge length. (Pfann, Zone Melting, 2d ed., Wiley, New
York, 1966, p. 290.)
FIG. 20-7 End-fed column crystallizer. (Phillips Petroleum Co.)
Figure 20-8 shows the features of a horizontal center-fed column
[Brodie, Aust. Mech. Chem. Eng. Trans., 37 (May 1979)] which has
been commercialized for continuous purification of naphthalene and
p-dichlorobenzene. Liquid feed enters the column between the hot
purifying section and the cold freezing or recovery zone. Crystals are
formed internally by indirect cooling of the melt through the walls of
the refining and recovery zones. Residue liquid that has been
depleted of product exits from the coldest section of the column. A
spiral conveyor controls the transport of solids through the unit.
Another center-fed design that has only been used on a preparative
scale is the vertical spiral conveyor column reported by Schildknecht
[Angew. Chem., 73, 612 (1961)]. In this device, a version of which is
shown on Fig. 20-9, the dispersed-crystal phase is formed in the freez-
ing section and conveyed downward in a controlled manner by a rotat-
ing spiral with or without a vertical oscillation.
Differences have been observed in the performance of end- and
center-fed column configurations. Consequently, discussions of center-
and end-fed column crystallizers are presented separately. The design
and operation of both columns are reviewed by Powers (Zief and
Wilcox, op. cit., p. 343). A comparison of these devices is shown on
Center-Fed Column Crystallizers Two types of center-fed col-
umn crystallizers are illustrated on Figs. 20-8 and 20-9. As in a simple
distillation column, these devices are composed of three distinct sec-
tions: a freezing or recovery section, where solute is frozen from the
impure liquor; the purification zone, where countercurrent contacting
of solids and liquid occurs; and the crystal-melting and -refluxing sec-
tion. Feed position separates the refining and recovery portions of the
purification zone. The section between feed location and melter is
referred to as the refining or enrichment section, whereas the section
between feed addition and freezing is called the recovery section. The
refining section may have provisions for sidewall cooling. The pub-
lished literature on column crystallizers connotes stripping and refin-
ing in a reverse sense to distillation terminology, since refined product
from a melt crystallizer exits at the hot section of the column rather
than at the cold end as in a distillation column.
Rate processes that describe the purification mechanisms in a col-
umn crystallizer are highly complex since phase transition and heat-
and mass-transfer processes occur simultaneously. Nucleation and
growth of a crystalline solid phase along with crystal washing and crys-
tal melting are occurring in various zones of the apparatus. Column
hydrodynamics are also difficult to describe. Liquid- and solid-phase
mixing patterns are influenced by factors such as solids-transport
mechanism, column orientation, and, particularly for dilute slurries,
the settling characteristics of the solids.
Most investigators have focused their attention on a differential
segment of the zone between the feed injection and the crystal melter.
Analysis of crystal formation and growth in the recovery section has
received scant attention. Table 20-4 summarizes the scope of the lit-
erature treatment for center-fed columns for both solid-solution and
eutectic forming systems.
The dominant mechanism of purification for column crystallization
of solid-solution systems is recrystallization. The rate of mass transfer
resulting from recrystallization is related to the concentrations of the
solid phase and free liquid which are in intimate contact. A model
based on height-of-transfer-unit (HTU) concepts representing the
composition profile in the purification section for the high-melting
component of a binary solid-solution system has been reported by
Powers et al. (in Zief and Wilcox, op. cit., p. 363) for total-reflux oper-
ation. Typical data for the purification of a solid-solution system,
azobenzene-stilbene, are shown in Fig. 20-10. The column crystallizer
was operated at total reflux. The solid line through the data was com-
puted by Powers et al. (op. cit., p. 364) by using an experimental HTU
value of 3.3 cm.
CRYSTALLIZATION FROM THE MELT 20-7
FIG. 20-9 Center-fed column crystallizer with a spiral-type conveyor.
TABLE 20-3 Comparison of Melt-Crystallizer Performance
Center-fed column End-fed column
Solid phase is formed internally; Solid phase is formed in external
thus, only liquid streams enter equipment and fed as slurry into
and exit the column. the purifier.
Internal reflux can be controlled The maximum internal liquid reflux
without affecting product yield. is fixed by the thermodynamic
state of the feed relative to the
product stream. Excessive reflux
will diminish product yield.
Operation can be continuous or Total reflux operation is not feasible.
batchwise at total reflux.
Center-fed columns can be adapted End-fed columns are inefficient
for both eutectic and solid- for separation of solid-solution
solution systems. systems.
Either low- or high-porosity- End-fed units are characterized by
solids-phase concentrations can low-porosity-solids packing in the
be formed in the purification purification and melting zones.
and melting zones.
Scale-up depends on the mechanical Scale-up is limited by design of
complexity of the crystal-transport melter and/or crystal-washing
system and techniques for removing section. Vertical or horizontal
heat. Vertical oscillating spiral columns of several meters in
columns are likely limited to about diameter are possible.
0.2 m in diameter, whereas
horizontal columns of several
meters are possible.
FIG. 20-8 Horizontal center-fed column crystallizer. (The C. W. Nofsinger Co.)
Most of the analytical treatments of center-fed columns describe
the purification mechanism in an adiabatic oscillating spiral column
(Fig. 20-9). However, the analyses by Moyers (op. cit.) and Griffin (op.
cit.) are for a nonadiabatic dense-bed column. Differential treatment
of the horizontal-purifier (Fig. 20-8) performance has not been
reported; however, overall material and enthalpy balances have been
described by Brodie (op. cit.) and apply equally well to other designs.
A dense-bed center-fed column (Fig. 20-11) having provision for
internal crystal formation and variable reflux was tested by Moyers
et al. (op. cit.). In the theoretical development (ibid.) a nonadiabatic,
plug-flow axial-dispersion model was employed to describe the per-
formance of the entire column. Terms describing interphase transport
of impurity between adhering and free liquid are not considered.
A comparison of the axial-dispersion coefficients obtained in oscil-
lating-spiral and dense-bed crystallizers is given in Table 20-5. The
dense-bed column approaches axial-dispersion coefficients similar to
those of densely packed ice-washing columns.
The concept of minimum reflux as related to column-crystallizer
operation is presented by Brodie (op. cit.) and is applicable to all types
of column crystallizers, including end-fed units. In order to stabilize
column operation the sensible heat of subcooled solids entering the
melting zone should be balanced or exceeded by the heat of fusion of
the refluxed melt. The relationship in Eq. (20-8) describes the mini-
mum reflux requirement for proper column operation.
R = (TP − TF) CP /λ (20-8)
R = reflux ratio, g reflux/g product; TP = product temperature, °C;
TF = saturated-feed temperature, °C; CP = specific heat of solid crys-
tals, cal/(g⋅°C); and λ = heat of fusion, cal/g.
All refluxed melt will refreeze if reflux supplied equals that com-
puted by Eq. (20-8). When reflux supplied is greater than the mini-
mum, jacket cooling in the refining zone or additional cooling in the
20-8 ALTERNATIVE SEPARATION PROCESSES
TABLE 20-4 Column-Crystallizer Investigations
Total reflux—steady state 1, 2, 4, 6 1, 4, 6
Total reflux—dynamic 2
Continuous—steady state 1, 4 4, 8, 9
Total reflux—steady state 1, 3, 4, 7 1, 3, 6
Continuous—steady state 1, 5, 10, 11, 12 5, 8, 9, 10, 11, 13
1. Powers, Symposium on Zone Melting and Column Crystallization, Karls-
2. Anikin, Dokl. Akad. Nauk SSSR, 151, 1139 (1969).
3. Albertins et al., Am. Inst. Chem. Eng. J., 15, 554 (1969).
4. Gates et al., Am. Inst. Chem. Eng. J., 16, 648 (1970).
5. Henry et al., Am. Inst. Chem. Eng. J., 16, 1055 (1970).
6. Schildknecht et al., Angew. Chem., 73, 612 (1961).
7. Arkenbout et al., Sep. Sci., 3, 501 (1968).
8. Betts et al., Appl. Chem., 17, 180 (1968).
9. McKay et al., Chem. Eng. Prog. Symp. Ser., no. 25, 55, 163 (1959).
10. Bolsaitis, Chem. Eng. Sci., 24, 1813 (1969).
11. Moyers et al., Am. Inst. Chem. Eng. J., 20, 1119 (1974).
12. Griffin, M.S. thesis in chemical engineering, University of Delaware, 1975.
13. Brodie, Aust. Mech. Chem. Eng. Trans., 37 (1971).
FIG. 20-10 Steady-state separation of azobenzene and stilbene in a center-fed
column crystallizer with total-reflux operation. To convert centimeters to inches,
multiply by 0.3937. (Zief and Wilcox, Fractional Solidification, vol. 1, Marcel
Dekker, New York, 1967, p. 356.)
FIG. 20-11 Dense-bed center-fed column crystallizer. [Moyers et al., Am.
Inst. Chem. Eng. J., 20, 1121 (1974).]
TABLE 20-5 Comparison of Axial-Dispersion Coefficients
for Several Liquid-Solid Contactors
Column type coefficient, cm2
Center-fed crystallizer (oscillating 1.6–3.5 1
Center-fed crystallizer (oscillating 1.3–1.7 2
Countercurrent ice-washing column 0.025–0.17 3
Center-fed crystallizer 0.12–0.30 4
1. Albertins et al., Am. Inst. Chem. Eng. J., 15, 554 (1969).
2. Gates et al., Am. Inst. Chem. Eng. J., 16, 648 (1970).
3. Ritter, Ph.D. thesis, Massachusetts Institute of Technology, 1969.
4. Moyers et al., Am. Inst. Chem. Eng. J., 20, 1119 (1974).
recovery zone is required to maintain product recovery. Since high-
purity melts are fed near their pure-component freezing tempera-
tures, little refreezing takes place unless jacket cooling is added.
To utilize a column-crystallizer design or rating model, a large
number of parameters must be identified. Many of these are empir-
ical in nature and must be determined experimentally in equipment
identical to the specific device being evaluated. Hence macroscopic
evaluation of systems by large-scale piloting is the rule rather than
the exception. Included in this rather long list of critical parameters
are factors such as impurity level trapped in the solid phase, prod-
uct quality as a function of reflux ratio, degree of liquid and solids
axial mixing in the equipment as a function of solids-conveyor
design, size and shape of crystals produced, and ease of solids han-
dling in the column. Heat is normally removed through metal sur-
faces; thus, the stability of the solution to subcooling can also be a
major factor in design.
End-Fed Column Crystallizer End-fed columns were devel-
oped and successfully commercialized by the Phillips Petroleum
Company in the 1950s. The sections of a typical end-fed column,
often referred to as a Phillips column, are shown on Fig. 20-7. Impure
liquor is removed through filters located between the product-
freezing zone and the melter rather than at the end of the freezing
zone, as occurs in center-fed units. The purification mechanism for
end-fed units is basically the same as for center-fed devices. However,
there are reflux restrictions in an end-fed column, and a high degree
of solids compaction exists near the melter of an end-fed device. It has
been observed that the free-liquid composition and the fraction of
solids are relatively constant throughout most of the purification sec-
tion but exhibit a sharp discontinuity near the melting section [McKay
et al., Ind. Eng. Chem., 52, 197 (1969)]. Investigators of end-fed col-
umn behavior are listed in Table 20-6. Note that end-fed columns are
adaptable only for eutectic-system purification and cannot be oper-
ated at total reflux.
Performance information for the purification of p-xylene indicates
that nearly 100 percent of the crystals in the feed stream are removed as
product. This suggests that the liquid which is refluxed from the melt-
ing section is effectively refrozen by the countercurrent stream of sub-
cooled crystals. A high-melting product of 99.0 to 99.8 weight percent
p-xylene has been obtained from a 65 weight percent p-xylene feed.
The major impurity was m-xylene. Figure 20-12 illustrates the column-
cross-section-area–capacity relationship for various product purities.
Column crystallizers of the end-fed type can be used for purifica-
tion of many eutectic-type systems and for aqueous as well as organic
systems (McKay, loc. cit.). Column crystallizers have been used for
xylene isomer separation, but recently other separation technologies
including more efficient melt crystallization equipment have tended
to supplant the Phillips style crystallizer.
Commercial Equipment and Applications In the last two
decades the practice of melt crystallization techniques for purification
of certain organic materials has made significant commercial progress.
The concept of refining certain products by countercurrent staging of
crystallization in a column has completed the transition from labora-
tory and pilot equipment to large-scale industrial configurations.
Chemicals which have been purified by suspension crystallization-
purifier column techniques are listed on Table 20-7. The practice of
crystal formation and growth from the bulk liquid (as is practiced in
suspension crystallization techniques described in Sec. 18 of this
handbook) and subsequent crystal melting and refluxing in a purifier
column has evolved into two slightly different concepts: (1) the hori-
zontal continuous crystallization technique with vertical purifier
invented by Brodie (op. cit.) and (2) the continuous multistage or
stepwise system with vertical purifier developed by Tsukishima Kikai
Co., Ltd. (TSK). A recent description of these processes has been
published by Meyer [Chem. Proc., 53, 50 (1990)].
The horizontal continuous Brodie melt crystallizer is basically an
indirectly cooled crystallizer with an internal ribbon conveyor to
transport crystals countercurrent to the liquid and a vertical purifier
for final refining. Figure 20-8 describes the operation of a single tube
unit and Fig. 20-13 depicts a multitube unit. The multitube design
has been successfully commercialized for a number of organic chem-
icals. The Brodie purifier configuration requires careful control of
process and equipment temperature differences to eliminate internal
encrustations and is limited by the inherent equipment geometry to
capacities of less than 15,000 tons per year per module.
In the multistage process described on Fig. 20-14 feed enters one
of several crystallizers installed in series. Crystals formed in each crys-
tallizer are transferred to a hotter stage and the liquid collected in the
clarified zone of the crystallizer is transferred to a colder stage and
eventually discharged as residue. At the hot end, crystals are trans-
ferred to a vertical purifier where countercurrent washing is performed
by pure, hot-product reflux. TSK refers to this multistage process as
the countercurrent cooling crystallization (CCCC) process. In
CRYSTALLIZATION FROM THE MELT 20-9
TABLE 20-6 End-Fed-Crystallizer Investigations
Eutectic systems Theoretical Experimental
Continuous—steady state 1, 2, 4 1, 4
Batch 3 3
1. McKay et al., Ind. Eng. Chem. Process Des. Dev., 6, 16 (1967).
2. Player, Ind. Eng. Chem. Process Des. Dev., 8, 210 (1969).
3. Yagi et al., Kagaku Kogaku, 72, 415 (1963).
4. Shen and Meyer, Prepr. 19F, AIChE Symp., Chicago, 1970.
TABLE 20-7 Chemicals Purified by TSK CCCC Process
(The C. W. Nofsinger Co.)
Dibutyl hydroxy toluene (BHT)
Hexachloro cyclo butene
Tertiary butyl phenol
FIG. 20-12 Pulsed-column capacity versus column size for 65 percent
p-xylene feed. To convert gallons per hour to cubic meters per hour, multiply by
0.9396; to convert square feet to square meters, multiply by 0.0929. (McKay
et al., prepr., 59th nat. meet. AIChE, East Columbus, Ohio.)
20-10 ALTERNATIVE SEPARATION PROCESSES
Refining section Recovery section
FIG. 20-13 Horizontal continuous Brodie melt crystallizer—multitube unit. (The C. W. Nofisinger Co.)
FIG. 20-14 Crystallizer—multistage process. (The C. W. Nofisinger Co.)
principle any suitable type of crystallizer can be used in the stages as
long as the crystals formed can be separated from the crystallizer liq-
uid and settled and melted in the purifier.
Commercial applications for both the Brodie and CCCC process
are indicated on Table 20-8. Both the Brodie Purifier and the CCCC
processes are available from The C. W. Nofsinger Company, PO Box
419173, Kansas City, MO 64141-0173.
Falling-film crystallization utilizes progressive freezing principles
to purify melts and solutions. The technique established to practice
the process is inherently cyclic. Figure 20-15 depicts the basic
working concept. First a crystalline layer is formed by subcooling a
liquid film on a vertical surface inside a tube. This coating is then
CRYSTALLIZATION FROM THE MELT 20-11
TABLE 20-8 Commercial TSK Crystallization Operating Plants
MM lb/yr Company Date & location
Countercurrent Cooling Crystallization (CCCC) Process
Nofsinger design & construct
p-Dichlorobenzene Confidential Monsanto Co. 1989—Sauget, IL
TSK design & construct
Confidential Confidential 1988—Japan
p-Xylene 137 MGC 1986—Mizushima, Japan
Confidential Confidential 1985—Japan
p-Xylene 132 MGC 1983—Mizushima, Japan
expanded to 160 1984—Japan
p-Xylene 26.5 MGC 1981—Mizushima, Japan
TSK design & construct
Naphthalene 8 SHSM 1985—China
p-DCB 13 Hodogaya 1981—Japan
p-DCB 5.5 Sumitomo 1978—Japan
UCAL license & design
Naphthalene 16 Nippon Steel 1974—Japan
p-DCB 13 Hodogaya 1974—Japan
UCAL license & design
Naphthalene 10 British Tar 1972—U.K.
p-DCB 3 UCAL 1969—Australia
Nofsinger The C. W. Nofsinger Company
TSK Tsukishima Kikai Co., Ltd.
SHSM Shanghai Hozan Steel Mill
UCAL Union Carbide Australia, Ltd.
MGC Mitsubishi Gas Chemical Co., co-developer with TSK of the application for p-Xylene
1. Commercial scale plant started up in the spring of 1988 purifying a bulk chemical. This is the first application of the
CCCC process on this bulk chemical.
2. This small unit is operating in Japan with an 800-mm crystallizer and 300-mm purifier. Because of confidentiality, we can-
not disclose the company, capacity, or product.
FIG. 20-15 Dynamic crystallization system. (Sulzer Chemtech.)
20-12 ALTERNATIVE SEPARATION PROCESSES
T-1 T-2 T-3
System flow sheet
P-4 P-3 P-5
FIG. 20-16 Sulzer MWB-crystallization process. (a) Stepwise operation of the process. (b) System flow sheet. (Sulzer Chemtech.)
10 20 30 40
50 60 70
80 90 100 110 120
2500 2500 1500
Heating medium temperature
CRYSTALLIZATION FROM THE MELT 20-13
grown by extracting heat from a falling film of melt (or solution)
through a heat transfer surface. Impure liquid is then drained from
the crystal layer and the product is reclaimed by melting. Variants
of this technique have been perfected and are used commercially
for many types of organic materials. Both static and falling-film
techniques have been described by Wynn [Chem. Eng. Progr.,
(1992)]. Mathematical models for both static and dynamic opera-
tions have been presented by Gilbert [AIChE J., 37, 1205 (1991)].
Principles of Operation Figure 20-16 describes a typical
three-stage falling-film crystallization process for purification of
MCA (monochloro acetic acid). Crystallizer E-8 consists of a num-
ber of vertical tubular elements working in parallel enclosed in a
shell. Normal tube length is 12 meters with a 50- to 75-millimeter
tube inside diameter. Feed enters stage two of the sequential oper-
ation, is added to the kettle (T-5), and is then circulated to the top
of the crystallizer and distributed as a falling film inside the tubes.
Nucleation is induced at the inside walls and a crystal layer starts to
grow. Temperature of the coolant is progressively lowered to com-
pensate for reduced heat transfer and lower melt freezing point
until the thickness inside the tube is between 5 and 20 millimeters
depending on the product. Kettle liquid is evactuated to the first-
stage holding tank (T-3) for eventual recrystallization at a lower
temperature to maximize product yield and to strip product from
the final liquid residue. Semirefined product frozen to the inside of
the tube during operation of stage two is first heated above its melt-
ing point and slightly melted (sweated). This semipurified melted
material (sweat) is removed from the crystallizer kettle, stored in a
stage tank (T-4), and then added to the next batch of fresh feed. The
remaining material inside the crystallizer is then melted, mixed
with product sweat from stage three, recrystallized, and sweated to
upgrade the purity even further (stage 3).
Commercial Equipment and Applications The falling-film
crystallization process was invented by the MWB company in Switzer-
land. The process is now marketed by Sulzer Chemtech. Products
successfully processed in the falling-film crystallizer are listed on
Table 20-9. The falling-film crystallization process is available from
the Chemtech Div. of Sulzer Canada Inc., 60 Worcester Rd., Rexdale,
Ontario N9W 5X2 Canada.
TABLE 20-9 Fractional Crystallization Reference List
Product Main characteristics tons/year Purity Type of plant Country Client
Acrylic acid Very low aldehyde content, Undisclosed 99.95% Falling film Undisclosed Undisclosed
no undesired polymerization Undisclosed 99.9% Falling film Undisclosed Undisclosed
in the plant
Benzoic acid Pharmaceutical grade, odor- and 4,500 99.97% Falling film Italy Chimica del Friuli
Bisphenol A Polycarbonate grade, no solvent 150,000 Undisclosed Falling film USA General Electric
Carbonic acid 1,200 Undisclosed Falling film Germany Undisclosed
Fatty acid Separation of tallow fatty acid 20,000 Stearic acid: Iodine no. 2 Falling film Japan Undisclosed
into saturated and Oleic acid: Cloud pt 5°C
Fine chemicals <1,000 Undisclosed Falling film GUS Undisclosed
<1,000 Undisclosed Static Switzerland Undisclosed
<1,000 Undisclosed Falling film Switzerland Undisclosed
<1,000 Undisclosed Falling film Switzerland Undisclosed
<1,000 Undisclosed Falling film USA Undisclosed
<1,000 Undisclosed Falling film Germany Undisclosed
<1,000 Undisclosed Falling film Japan Undisclosed
Hydrazine Satellite grade 3 >99.9% Falling film Germany ESA
Monochloro acetic Low DCA content 6,000 >99.2% Falling film USA Undisclosed
Multipurpose Separation or purification of two 1,000 Various grades Falling film Belgium UCB
or more chemicals, alternatively 1,000 Undisclosed Falling film Belgium Reibelco
Naphthalene Color free and color stable with 60,000 99.5% Falling film Germany Rütgers-Werke
low thionaphthene content 20,000 99.5% Falling film/static P.R. China Anshan
10,000 99.8% Falling film/static P.R. China Jining
12,000 Various grades Falling film The Netherlands Cindu Chemicals
p-Dichlorobenzene No solvent washing required 40,000 99.95% Falling film USA Standard Chlorine
5,000 99.98% Falling film Japan Toa Gosei
4,000 99.8% Falling film/distillation Brazil Nitroclor
3,000 99.95% Falling film P.R. China Fuyang
3,000 >97% Static P.R. China Shandong
p-Nitrochlorobenzene 18,000 99.3% Falling film/distillation P.R. China Jilin Chemical
10,000 99.5% Static India Mardia
Toluene diisocyanate Separation of TDI 80 into 20,000 Undisclosed Falling film Undisclosed Undisclosed
(TDI) TDI 100 & TDI 65
Trioxan <1,000 99.97% Falling film Undisclosed Undisclosed
GENERAL REFERENCES: Yeo and Kiran, J. Supercritical Fluids, 34, 287–308
(2005). York, Kompella, and Shekunov, Supercritical Fluid Technology for Drug
Product Development, Marcel Dekker, New York, 2004. Shah, Hanrath, Johnston,
and Korgel, J. Physical Chemistry B, 108, 9574–9587 (2004). Eckert, Liotta,
Bush, Brown, and Hallett, J. Physical Chemistry B, 108, 18108–18118 (2004).
DeSimone, Science, 297, 799–803 (2002). Arai, Sako, and Takebayashi, Super-
critical Fluids: Molecular Interactions, Physical Properties, and New Applications,
Springer, New York, 2002. Kiran, Debenedetti, and Peters, Supercritical Fluids:
Fundamentals and Applications, Kluwer Academic, Dordrecht, 2000. McHugh
and Krukonis, Supercritical Fluid Extraction Principles and Practice, 2d ed., But-
terworth, Stoneham, Mass., 1994. Brunner, Gas Extraction: An Introduction to
Fundamentals of Supercritical Fluids and the Application to Separation Processes,
Springer, New York, 1994. Gupta and Shim, Solubility in Supercritical Carbon
Dioxide, CRC Press, Boca Raton, Fla., 2007. Gupta and Kompella, Nanoparticle
Technology for Drug Delivery, Taylor & Francis, New York, 2006.
Fluids above their critical temperatures and pressures, called super-
critical fluids (SCFs), exhibit properties intermediate between those
of gases and liquids. Consequently, each of these two boundary condi-
tions sheds insight into the nature of these fluids. Unlike gases, SCFs
possess a considerable solvent strength, and transport properties are
more favorable. In regions where a SCF is highly compressible, its
density and hence its solvent strength may be adjusted over a wide
range with modest variations in temperature and pressure. This tun-
ability may be used to control phase behavior, separation processes
(e.g., SCF extraction), rates and selectivities of chemical reactions,
and morphologies in materials processing. For SCF separation
processes to be feasible, the advantages (Table 20-10) must compen-
sate for the costs of high pressure; examples of commercial applica-
tions are listed in Table 20-11.
The two SCFs most often studied—CO2 and water—are the two
least expensive of all solvents. CO2 is nontoxic and nonflammable
and has a near-ambient critical temperature of 31.1°C. CO2 is an
environmentally friendly substitute for organic solvents including
chlorocarbons and chlorofluorocarbons. Supercritical water (Tc =
374°C) is of interest as a substitute for organic solvents to minimize
waste in extraction and reaction processes. Additionally, it is used for
hydrothermal oxidation of hazardous organic wastes (also called
supercritical water oxidation) and hydrothermal synthesis. (See also
Sec. 15 for additional discussion of supercritical fluid separation
PHYSICAL PROPERTIES OF PURE SUPERCRITICAL FLUIDS
Thermodynamic Properties The variation in solvent strength
of a SCF from gaslike to liquidlike values (see Table 20-12) may be
described qualitatively in terms of the density ρ, as shown in Fig.
20-17, or the solubility parameter.
Similar characteristics are observed for other density-dependent
variables including enthalpy, entropy, viscosity, and diffusion coeffi-
cient. Above the critical temperature, it is possible to tune the solvent
20-14 ALTERNATIVE SEPARATION PROCESSES
TABLE 20-10 Advantages of Supercritical Fluid Separations
Solvent strength is adjustable to tailor selectivities and yields.
Diffusion coefficients are higher and viscosities lower, compared with liquids.
Low surface tension favors wetting and penetration of small pores.
There is rapid diffusion of CO2 through condensed phases, e.g., polymers
and ionic liquids.
Solvent recovery is fast and complete, with minimal residue in product.
Collapse of structure due to capillary forces is prevented during solvent
Properties of CO2 as a solvent:
Environmentally acceptable solvent for waste minimization, nontoxic,
nonflammable, inexpensive, usable at mild temperatures.
Properties of water as a solvent:
Nontoxic, nonflammable, inexpensive substitute for organic solvents.
Extremely wide variation in solvent strength with temperature and pressure.
TABLE 20-11 Commercial Applications of Supercritical Fluid
Extraction of foods and pharmaceuticals
Caffeine from coffee and tea
Flavors, cholesterol, and fat from foods
Nicotine from tobacco
Solvents from pharmaceutical compounds and drugs from natural sources
Extraction of volatile substances from substrates
Drying and aerogel formation
Cleaning fabrics, quartz rods for light guide fibers, residues in microelectronics
Removal of monomers, oligomers, and solvent from polymers
Residuum oil supercritical extraction (ROSE) (petroleum deasphalting)
Polymer and edible oils fractionation
CO2 enhanced oil recovery
Analytical SCF extraction and chromatography
Infusion of materials into polymers (dyes, pharmaceuticals)
Extraction of sec-butanol from isobutene
Polymerization to form Teflon
Depolymerization, e.g., polyethylene terephthalate and cellulose hydrolysis
Hydrothermal oxidation of organic wastes in water
Crystallization, particle formation, and coatings
Antisolvent crystallization, rapid expansion from supercritical fluid solution
Particles from gas saturated solutions
Crystallization by reaction to form metals, semiconductors (e.g., Si), and
metal oxides including nanocrystals
Supercritical fluid deposition
SUPERCRITICAL FLUID SEPARATION PROCESSES
TABLE 20-12 Physical Properties of a Supercritical Fluid Fall
between Those of a Typical Gas and Liquid
Liquid Supercritical fluid Gas
Density, g/mL 1 0.05 – 1 10−3
Viscosity, Pa·s 10−3
Diffusion coefficients, cm2
Surface tension, mN/m 20–50 0 0
FIG. 20-17 Density versus pressure and temperature for CO2. (Tc = 31.1°C,
Pc = 73.8 bar.)
0 50 100 150 200 250 300 350
strength continuously over a wide range with either a small isothermal
pressure change or a small isobaric temperature change. This unique
ability to tune the solvent strength of a SCF may be used to extract
and then recover selected products. A good indicator of the van der
Waals forces contributed by a SCF is obtained by multiplying ρ by the
molecular polarizability α, which is a constant for a given molecule.
CO2’s small αρ and low solvent strength are more like those of a fluo-
rocarbon than those of a hydrocarbon.
Water, a key SCF, undergoes profound changes upon heating to the
critical point. It expands by a factor of 3, losing about two-thirds of
the hydrogen bonds, the dielectric constant drops from 80 to 5 (Shaw
et al., op. cit.), and the ionic product falls several orders or magnitude
(see Fig. 20-18). At lower densities, supercritical water (SCW) behaves
as a “nonaqueous” solvent, and it dissolves many organics and even
gases such as O2. Here it does not solvate ions significantly.
Transport Properties Although the densities of SCFs can
approach those of conventional liquids, transport properties are more
favorable because viscosities remain lower and diffusion coefficients
remain higher. Furthermore, CO2 diffuses through condensed-liquid
phases (e.g., adsorbents and polymers) faster than do typical solvents
which have larger molecular sizes. For example, at 35o
C the estimated
pyrene diffusion coefficient in polymethylmethacrylate increases by
4 orders of magnitude when the CO2 content is increased from 8 to
17 wt % with pressure [Cao, Johnston, and Webber, Macromolecules,
38(4), 1335–1340 (2005)].
Liquid-Fluid Equilibria Nearly all binary liquid-fluid phase
diagrams can be conveniently placed in one of six classes (Prausnitz,
Lichtenthaler, and de Azevedo, Molecular Thermodynamics of Fluid
Phase Equilibria, 3d ed., Prentice-Hall, Upper Saddle River, N.J., 1998).
Two-phase regions are represented by an area and three-phase regions
by a line. In class I, the two components are completely miscible, and a
single critical mixture curve connects their critical points. Other classes
may include intersections between three phase lines and critical curves.
For a ternary system, the slopes of the tie lines (distribution coefficients)
and the size of the two-phase region can vary significantly with pressure
as well as temperature due to the compressibility of the solvent.
Solid-Fluid Equilibria The solubility of the solid is very sensi-
tive to pressure and temperature in compressible regions, where the
solvent’s density and solubility parameter are highly variable. In con-
trast, plots of the log of the solubility versus density at constant tem-
perature often exhibit fairly simple linear behavior (Fig. 20-19). To
understand the role of solute-solvent interactions on solubilities and
selectivities, it is instructive to define an enhancement factor E as
the actual solubility y2 divided by the solubility in an ideal gas, so that
E = y2P/P2
, where P2
is the vapor pressure. The solubilities in CO2
are governed primarily by vapor pressures, a property of the solid
crystals, and only secondarily by solute-solvent interactions in the
SCF phase. For example, for a given fluid at a particular T and P, the
E’s are similar for the three sterols, each containing one hydroxyl
group, even though the actual solubilities vary by many orders of
magnitude (Fig. 20-19).
Polymer-Fluid Equilibria and the Glass Transition Most
polymers are insoluble in CO2, yet CO2 can be quite soluble in the
polymer-rich phase. The solubility in CO2 may be increased by a
combination of lowering the cohesive energy density (which is pro-
portional to the surface tension of the polymer [O’Neill et al., Ind.
Eng. Chem. Res., 37, 3067–79 (1998)]), branching, and the incor-
poration of either acetate groups in the side chain or carbonate
groups in the backbone of the polymer [Sarbu, Styranec, and Beck-
man, Nature, 405, 165–168 (2000)]. Polyfluoromethacrylates are
extremely soluble, and functionalized polyethers and copolymers of
cyclic ethers and CO2 have been shown to be more soluble than
most other nonfluorinated polymers. The sorption of CO2 into sili-
cone rubber is highly dependent upon temperature and pressure,
since these properties have a large effect on the density and activity
of CO2. For glassy polymers, sorption isotherms are more complex,
and hysteresis between the pressurization and depressurization
steps may appear. Furthermore, CO2 can act as a plasticizer and
depress the glass transition temperature by as much as 100°C or
even more, producing large changes in mechanical properties and
diffusion coefficients. This phenomenon is of interest in condition-
ing membranes for separations and in commercial foaming of poly-
mers to reduce VOC emissions.
Cosolvents and Complexing Agents Many nonvolatile polar
substances cannot be dissolved at moderate temperatures in nonpolar
fluids such as CO2. Cosolvents (also called entrainers) such as alcohols
and acetone have been added to fluids to raise the solvent strength for
organic solutes and even metals. The addition of only 2 mol % of the
complexing agent tri-n-butyl phosphate (TBP) to CO2 increases the
solubility of hydroquinone by a factor of 250 due to Lewis acid-base
Surfactants and Colloids in Supercritical Fluids Because
very few nonvolatile molecules are soluble in CO2, many types of
hydrophilic or lipophilic species may be dispersed in the form of
polymer latexes (e.g., polystyrene), microemulsions, macroemul-
sions, and inorganic suspensions of metals and metal oxides (Shah
et al., op. cit.). The environmentally benign, nontoxic, and nonflam-
mable fluids water and CO2 are the two most abundant and inex-
pensive solvents on earth. Fluorocarbon and hydrocarbon-based
surfactants have been used to form reverse micelles, water-in-CO2
SUPERCRITICAL FLUID SEPARATION PROCESSES 20-15
FIG. 20-18 Physical properties of water versus temperature at 240 bar.
[Reprinted from Kritzer and Dinjus, “An Assessment of Supercritical Water Oxi-
dation (SCWO): Existing Problems, Possible Solutions and New Reactor Con-
cepts,” Chem. Eng. J., vol. 83(3), pp. 207–214, copyright 2001, with permission
100 200 300
400 500 600
FIG. 20-19 Solubility of sterols in pure CO2 at 35°C [Wong and Johnston,
Biotech. Prog., 2, 29 (1986)].
16 18 20 22
microemulsions (2- to 10-nm droplets) and macroemulsions (50-nm
to 5-µm droplets) in SCFs including CO2. These organized molecular
assemblies extend SCF technology to include nonvolatile hydrophilic
solutes and ionic species such as amino acids and even proteins. Sur-
factant micelles or microemulsions are used commercially in dry
cleaning and have been proposed for applications including polymer-
ization, formation of inorganic and pharmaceutical particles, and
removal of etch/ash residues from low-k dielectrics used in microelec-
tronics. CO2-in-water macroemulsions, stabilized by surfactants with
the proper hydrophilic-CO2-philic balance, are used in enhanced oil
recovery to raise the viscosity of the flowing CO2 phase for mobility
control. Alkane ligands with various head groups have been used to
stabilize inorganic nanocrystals in SCW and to stabilize Si and Ge
nanocrystals in SCF hydrocarbons and CO2 at temperatures from
350 to 500°C. Furthermore, colloids may be stabilized by electrostatic
stabilization in CO2 [Smith, Ryoo, and Johnston, J. Phys. Chem. B.,
109(43), 20155 (2005)].
Phase Equilibria Models Two approaches are available for
modeling the fugacity of a solute in a SCF solution. The compressed
gas approach includes a fugacity coefficient which goes to unity for an
ideal gas. The expanded liquid approach is given as
= xiγi(P°, xi) fi°L
where xi is the mole fraction, γi is the activity coefficient, P° and fi°
are the reference pressure and fugacity, respectively, and v⎯
i is the
partial molar volume of component i. A variety of equations of state
have been applied in each approach, ranging from simple cubic
equations such as the Peng-Robinson equation of state to the more
complex statistical associating fluid theory (SAFT) (Prausnitz et al.,
op. cit.). SAFT is successful in describing how changes in H bond-
ing of SCF water influence thermodynamic and spectroscopic
Experimental gas-solid mass-transfer data have been obtained for
naphthalene in CO2 to develop correlations for mass-transfer coeffi-
cients [Lim, Holder, and Shah, Am. Chem. Soc. Symp. Ser., 406, 379
(1989)]. The mass-transfer coefficient increases dramatically near the
critical point, goes through a maximum, and then decreases gradually.
The strong natural convection at SCF conditions leads to higher mass-
transfer rates than in liquid solvents. A comprehensive mass-transfer
model has been developed for SCF extraction from an aqueous phase
to CO2 in countercurrent columns [Seibert and Moosberg, Sep. Sci.
Technol., 23, 2049 (1988); Brunner, op. cit.].
PROCESS CONCEPTS IN SUPERCRITICAL
Figure 20-20 shows a one-stage extraction process that utilizes the
adjustability of the solvent strength with pressure or temperature.
The solvent flows through the extraction chamber at a relatively
high pressure to extract the components of interest from the feed.
The products are then recovered in the separator by depressuriza-
tion, and the solvent is recompressed and recycled. The products
can also be precipitated from the extract phase by raising the tem-
perature after the extraction to lower the solvent density. Multiple
extractions or multiple stages may be used with various profiles,
e.g., successive increases in pressure or decreases in pressure.
Solids may be processed continuously or semicontinuously by
pumping slurries or by using lock hoppers. For liquid feeds, multi-
stage separation may be achieved by continuous countercurrent
extraction, much as in conventional liquid-liquid extraction. In
SCF chromatography, selectivity may be tuned with pressure and
temperature programming, with greater numbers of theoretical
stages than in liquid chromatography and lower temperatures than
in gas chromatography.
Decaffeination of Coffee and Tea This application is driven
by the environmental acceptability and nontoxicity of CO2 as well as
by the ability to tailor the extraction with the adjustable solvent
strength. It has been practiced industrially for more than two
decades. Caffeine may be extracted from green coffee beans, and the
aroma is developed later by roasting. Various methods have been pro-
posed for recovery of the caffeine, including washing with water and
Extraction of Flavors, Fragrances, Nutraceuticals, and
Pharmaceuticals Flavors and fragrances extracted by using super-
critical CO2 are significantly different from those extracted by using
steam distillation or solvent extraction. The SCF extract can almost be
viewed as a new product due to changes in composition associated
with the greater amounts of extraction, as shown in Table 20-13. In
many instances the flavor or fragrance of the extract obtained with
CO2 is closer to the natural one relative to steam distillation.
Temperature-Controlled Residuum Oil Supercritical
Extraction (ROSE) The Kerr-McGee ROSE process has been
used worldwide for over two decades to remove asphaltenes from oil.
The extraction step uses a liquid solvent that is recovered at supercrit-
ical conditions to save energy, as shown in Fig. 20-21. The residuum is
contacted with butane or pentane to precipitate the heavy asphaltene
fraction. The extract is then passed through a series of heaters, where
it goes from the liquid state to a lower-density SCF state. Because the
entire process is carried out at conditions near the critical point, a rel-
atively small temperature change is required to produce a fairly large
density change. After the light oils have been removed, the solvent is
cooled back to the liquid state and recycled.
Polymer Devolatilization, Fractionation, and Plasticization
Supercritical fluids may be used to extract solvent, monomers, and
oligomers from polymers, including biomaterials. After extraction the
pressure is reduced to atmospheric, leaving little residue in the
20-16 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-20 Idealized diagram of a supercritical fluid extraction process for
Solid (2) rich SCF
3, 4, ...
2, 3, 4, ...
Expansion or heating
TABLE 20-13 Comparison of Percent Yields of Flavors
and Fragrances from Various Natural Products*
Steam distillation Supercritical CO2
Natural substance (% yield) (% yield)
Ginger 1.5–3.0 4.6
Garlic 0.06–0.4 4.6
Pepper 1.0–2.6 8–18
Rosemary 0.5–1.1 7.5
*Mukhopadhyay, Natural Extracts Using Supercritical Carbon Dioxide, CRC
Press, Boca Raton, Fla., 2000; Moyler, Extraction of flavours and fragrances
with compressed CO2, in Extraction of Natural Products Using Near-Critical
Solvents, King and Bott (eds.), Blackie Academic & Professional, London, 1993.
substrate; furthermore, the extracted impurities are easily recovered
from the SCF. The swelling and lowering of the glass transition tem-
perature of the polymer by the SCF can increase mass-transfer rates
markedly. This approach was used to plasticize block copolymer tem-
plates for the infusion of reaction precursors in the synthesis of
porous low-k dielectrics. For homopolymers, plasticization may be
used to infuse dyes, pharmaceuticals, etc., and then the SCF may be
removed to trap the solute in the polymer matrix. SCFs may be used
to fractionate polymers on the basis of molecular weight and/or com-
position with various methods for programming pressure and/or tem-
perature (McHugh, op. cit.).
Drying and Aerogel Formation One of the oldest applications
of SCF technology, developed in 1932, is SCF drying. The solvent is
extracted from a porous solid with a SCF; then the fluid is depressur-
ized. Because the fluid expands from the solid without crossing a liquid-
vapor phase transition, capillary forces that would collapse the structure
are not present. Using SCF drying, aerogels have been prepared with
densities so low that they essentially float in air and look like a cloud of
smoke. Also, the process is used in a commercial instrument to dry sam-
ples for electron microscopy without perturbing the structure.
Cleaning SCFs such as CO2 can be used to clean and degrease
quartz rods utilized to produce optical fibers, products employed in
the fabrication of printed-circuit boards, oily chips from machining
operations, precision bearings in military applications, and so on.
Research is in progress for removing residues in etch/ash processes in
Microelectronics Processing SCF CO2 is proposed as a “dry,”
environmentally benign processing fluid enabling replacement of
aqueous and organic solvents in microelectronics fabrication (Desi-
mone, op. cit.). Proposed applications include drying, lithography, sol-
vent spin coating, stripping, cleaning, metal deposition, and chemical
mechanical planarization. Due to its low surface tension, tunable sol-
vent strength, and excellent mass-transfer properties, CO2 offers
advantages in wetting of surfaces and small pores, and in removal of
contaminants at moderate temperatures.
Precipitation/Crystallization to Produce Nano- and Micropar-
ticles Because fluids such as CO2 are weak solvents for many solutes,
they are often effective antisolvents in fractionation and precipitation.
In general, a fluid antisolvent may be a compressed gas, a gas-expanded
liquid, or a SCF. Typically a liquid solution is sprayed through a nozzle
into CO2 to precipitate a solute. As CO2 mixes with the liquid phase, it
decreases the cohesive energy density (solvent strength) substantially,
leading to precipitation of dissolved solutes (e.g., crystals of proges-
terone). The high diffusion rates of the organic solvent into CO2 and
vice versa can lead to rapid phase separation, and the supersaturation
curve may be manipulated to vary the crystalline morphology (Yeo and
Kiran, op. cit.).
Nanoparticles of controllable size can be obtained in the supercrit-
ical antisolvent-enhanced mass-transfer (SAS-EM) process, which can
produce commercial quantities of pharmaceuticals (see Fig. 20-22)
[Chattopadhyay and Gupta, Ind. Engr. Chem. Res., 40, 3530– 3539
(2001)]. Here, the solution jet is injected onto an ultrasonic vibrating
surface H inside the antisolvent chamber to aid droplet atomization.
The particle size is controlled by varying the vibration intensity. For
most pharmaceuticals, organic compounds, proteins, and polymers,
average particle diameters range from 100 to 1000 nm; even smaller
particles may be obtained for certain inorganic compounds.
Rapid Expansion from Supercritical Solution and Particles
from Gas Saturated Solutions Rapid expansion from supercriti-
cal solution (RESS) of soluble materials may be used to form
microparticles or microfibers. A variety of inorganic crystals have
been formed naturally and synthetically in SCF water, and organic
crystals have been formed in SCF CO2. Recently, the addition of a
solid cosolvent (e.g., menthol, which can be removed later by subli-
mation) has overcome key limitations by greatly enhancing solubil-
ities in CO2 and producing smaller nanoparticles by reducing
particle-particle coagulation [Thakur and Gupta, J. Sup. Fluids, 37,
307–315 (2006)]. Another approach is to expand the solutions into
aqueous solutions containing soluble surfactants to arrest growth
due to particle collisions. RESS typically uses dilute solutions. For
concentrated solutions, the process is typically referred to as parti-
cle formation from gas saturated solutions (PGSS). Here CO2 low-
ers the viscosity of the melt to facilitate flow. Union Carbide
developed the commercial UNICARB process to replace organic
solvents with CO2 as a diluent in coating applications to reduce
volatile organic carbon emissions and form superior coatings. For
aqueous solutions, the expansion of CO2 facilitates atomization, and
the resulting cooling may be used to control the freezing of the
Reactive Separations Reactions may be integrated with SCF
separation processes to achieve a large degree of control for produc-
ing a highly purified product. Reaction products may be recovered by
SUPERCRITICAL FLUID SEPARATION PROCESSES 20-17
FIG. 20-21 Schematic diagram of the Kerr-McGee ROSE process.
volatilization into, or precipitation from, a SCF phase. A classic exam-
ple is the high-pressure production of polyethylene in SCF ethylene.
The molecular weight distribution may be controlled by choosing the
temperature and pressure for precipitating the polymer from the SCF
phase. Over a decade ago, Idemitsu commercialized a 5000 metric ton
per/year (t/yr) integrated reaction and separation process in SCF
isobutene. The reaction of isobutene and water produces sec-butanol,
which is extracted from water by the SCF solvent. SCF solvents have
been tested for reactive extractions of liquid and gaseous fuels from
heavy oils, coal, oil shale, and biomass. In some cases the solvent par-
ticipates in the reaction, as in the hydrolysis of coal and heavy oils with
SCW. Related applications include conversion of cellulose to glucose
in water, delignification of wood with ammonia, and liquefaction of
lignin in water.
Gas-expanded liquids (GXLs) are emerging solvents for environ-
mentally benign reactive separation (Eckert et al., op. cit.). GXLs,
obtained by mixing supercritical CO2 with normal liquids, show inter-
mediate properties between normal liquids and SCFs both in solvation
power and in transport properties; and these properties are highly tun-
able by simple pressure variations. Applications include chemical reac-
tions with improved transport, catalyst recycling, and product
Hydrothermal oxidation (HO) [also called supercritical water oxida-
tion (SCWO)] is a reactive process to convert aqueous wastes to water,
CO2, O2, nitrogen, salts, and other by-products. It is an enclosed and
complete water treatment process, making it more desirable to the
public than incineration. Oxidation is rapid and efficient in this one-
phase solution, so that wastewater containing 1 to 20 wt % organics
may be oxidized rapidly in SCW with the potential for higher energy
efficiency and less air pollution than in conventional incineration.
Temperatures range from about 375 to 650°C and pressures from
3000 to about 5000 psia.
Crystallization by Chemical Reaction
Supercritical Fluid Deposition (SFD) Metal films may be
grown from precursors that are soluble in CO2. The SFD process
yields copper films with fewer defects than those possible by using
chemical vapor deposition, because increased precursor solubility
removes mass-transfer limitations and low surface tension favors pene-
tration of high-aspect-ratio features [Blackburn et al., Science, 294,
High-Temperature Crystallization The size-tunable optical
and electronic properties of semiconductor nanocrystals are attrac-
tive for a variety of optoelectronic applications. In solution-phase
crystallization, precursors undergo chemical reaction to form
nuclei, and particle growth is arrested with capping ligands that
passivate the surface. However, temperatures above 350°C are typ-
ically needed to crystallize the group IV elements silicon and ger-
manium, due to the covalent network structure. Whereas liquid
solvents boil away at these elevated temperatures, SCFs under
pressure are capable of solvating the capping ligands to stabilize the
nanocrystals (Shah et al., op. cit.). Crystalline Si and Ge nanocrys-
tals, with an average size of 2 to 70 nm, may be synthesized in
supercritical CO2, hexane, or octanol at 400 to 550°C and 20 MPa in
a simple continuous flow reactor. UV-visible absorbance and photo-
luminescence (PL) spectra of Ge nanocrystals of 3- to 4-nm diame-
ter exhibit optical absorbance and PL spectra blue-shifted by
approximately 1.7 eV relative to the band gap of bulk Ge, as shown
in Fig. 20-23. One-dimensional silicon nanowires may be grown
from relatively monodisperse gold nanocrystals stabilized with
dodecanethiol ligands, as shown in Fig. 20-24. The first crystalline
silicon nanowires with diameters smaller than 5 nm and lengths
greater than 1 µm made by any technique were produced in SCF
hexane. Hydrothermal crystallization has also been used to produce
metal oxide nano- and microcrystals by rapid generation of super-
saturation during hydrolysis of precursors, such as metal nitrates,
during rapid heating of aqueous solutions.
20-18 ALTERNATIVE SEPARATION PROCESSES
300 400 500 600 700 800
<d> = 4.2 nm
<d> = 3.1 nm
FIG. 20-23 Normalized photoluminescence spectra of 3.1-nm (λexcitation = 320
nm) and 4.2-nm (λexcitation = 340 nm) Ge nanoparticles dispersed in chloroform at
25ЊC with quantum yields of 6.6 and 4.6 percent, respectively. [Reprinted with
permission from Lu et al., Nano Lett., 4(5), 969–974 (2004). Copyright 2004
American Chemical Society.]
FIG. 20-22 Schematic of supercritical antisolvent with enhanced mass-transfer process to produce nanoparticles of controllable size. R, pre-
cipitation chamber; SCF pump, supply of supercritical CO2; I, inline filter; H, ultrasonic horn; P, pump for drug solution; G, pressure gauge.
ALTERNATIVE SOLID/LIQUID SEPARATIONS 20-19
SEPARATION PROCESSES BASED PRIMARILY
ON ACTION IN AN ELECTRIC FIELD
Differences in mobilities of ions, molecules, or particles in an electric
field can be exploited to perform useful separations. Primary empha-
sis is placed on electrophoresis and dielectrophoresis. Analogous sep-
aration processes involving magnetic and centrifugal force fields are
widely applied in the process industry (see Secs. 18 and 19).
Theory of Electrical Separations
GENERAL REFERENCES: Newman, Adv. Electrochem. Electrochem. Eng., 5,
87 (1967); Ind. Eng. Chem., 60(4), 12 (1968). Ptasinski and Kerkhof, Sep. Sci.
Technol., 27, 995 (1992).
For electrolytic solutions, migration of charged species in an electric
field constitutes an additional mechanism of mass transfer. Thus the
flux of an ionic species Ni in (g⋅mol)/(cm2
⋅s) in dilute solutions can be
Ni = −zi uiᏲci∇E − Di∇ci + ci v (20-10)
The ionic mobility ui is the average velocity imparted to the species
under the action of a unit force (per mole). v is the stream velocity, cm/s.
In the present case, the electrical force is given by the product of the
electric field ∇E in V/cm and the charge ziᏲ per mole, where Ᏺ is the
Faraday constant in C/g equivalent and zi is the valence of the ith species.
Multiplication of this force by the mobility and the concentration ci
] yields the contribution of migration to the flux of the ith
The diffusive and convective terms in Eq. (20-10) are the same as in
nonelectrolytic mass transfer. The ionic mobility ui, (g⋅mol⋅cm2
can be related to the ionic-diffusion coefficient Di, cm2
/s, and the ionic
conductance of the ith species λi, cm2
ui = Di /RT = λi /|zi|Ᏺ2
where T is the absolute temperature, K; and R is the gas constant,
8.3143 J/(K⋅mol). Ionic conductances are tabulated in the literature
(Robinson and Stokes, Electrolyte Solutions, Academic, New York,
1959). For practical purposes, a bulk electrolytic solution is electri-
zi ci = 0 (20-12)
since the forces required to effect an appreciable separation of charge
are prohibitively large.
The current density (A/cm2
) produced by movement of charged
species is described by summing the terms in Eq. (20-13) for all
i = Ᏺ Αi
ziNi = −κ ∇E − Ᏺ Αi
ziDi ∇ci (20-13)
where the electrical conductivity κ in S/cm is given by
κ = Ᏺ2
ui ci (20-14)
In solutions of uniform composition, the diffusional terms vanish and
Eq. (20-13) reduces to Ohm’s law.
Conservation of each species is expressed by the relation
∂ci /∂t = −∇ ⋅ Ni (20-15)
provided that the species is not produced or consumed in homoge-
neous chemical reactions. In two important cases, this conservation
law reduces to the equation of convective diffusion:
(∂ci /∂t) + v∇ ⋅ ci = D ∇2
First, when a large excess of inert electrolyte is present, the electric field
will be small and migration can be neglected for minor ionic compo-
nents; Eq. (20-16) then applies to these minor components, where D is
the ionic-diffusion coefficient. Second, Eq. (20-16) applies when the
solution contains only one cationic and one anionic species. The electric
field can be eliminated by means of the electroneutrality relation.
In the latter case the diffusion coefficient D of the electrolyte is
D = (z+ u+ D− − z− u− D+)/(z+ u+ − z− u− ) (20-17)
which represents a compromise between the diffusion coefficients of
the two ions. When Eq. (20-16) applies, many solutions can be ob-
tained by analogy with heat transfer and nonelectrolytic mass transfer.
Because the solution is electrically neutral, conservation of charge
is expressed by differentiating Eq. (20-13):
∇ ⋅ i = 0 = −κ ∇2
E − Ᏺ Αi
zi Di ∇2
For solutions of uniform composition, Eq. (20-18) reduces to
Laplace’s equation for the potential:
E = 0 (20-19)
This equation is the starting point for determination of the current-
density distributions in many electrochemical cells.
FIG. 20-24 High-resolution TEM image of Si nanowires produced at 500ºC and 24.1 MPa in supercritical hexane from gold seed crystals.
Inset: Electron diffraction pattern indexed for the <111> zone axis of Si indicates <110> growth direction. [Reprinted with permission from Lu
et al., Nano Lett., 3(1), 93–99 (2003). Copyright 2003 American Chemical Society.]
ALTERNATIVE SOLID/LIQUID SEPARATIONS
Near an interface or at solution junctions, the solution departs from
electroneutrality. Charges of one sign may be preferentially adsorbed
at the interface, or the interface may be charged. In either case, the
charge at the interface is counterbalanced by an equal and opposite
charge composed of ions in the solution. Thermal motion prevents
this countercharge from lying immediately adjacent to the interface,
and the result is a “diffuse-charge layer” whose thickness is on the
order of 10 to 100 Å.
A tangential electric field ∇Et acting on these charges produces a
relative motion between the interface and the solution just outside the
diffuse layer. In view of the thinness of the diffuse layer, a balance of
the tangential viscous and electrical forces can be written
) + ρe ∇Et = 0 (20-20)
where µ is the viscosity and ρe is the electric-charge density, C/cm3
Furthermore, the variation of potential with the normal distance sat-
isfies Poisson’s equation:
= −(ρe /ε) (20-21)
with ε defined as the permittivity of the solution. [The relative
dielectric constant is ε/ε0, where ε0 is the permittivity of free space;
ε0 = 8.8542 × 10−14
C/(V⋅cm).] Elimination of the electric-charge den-
sity between Eqs. (20-20) and (20-21) with two integrations, gives a
relation between ∇Et and the velocity v0 of the bulk solution relative
to the interface.
µ[vt(∞) − vt(0)] = ε ∇Et [E(∞) − E(0)] (20-22)
or v0 = −(ε ∇Etζ/µ) (20-23)
The potential difference across the mobile part of the diffuse-charge
layer is frequently called the zeta potential, ζ = E(0) − E(∞). Its value
depends on the composition of the electrolytic solution as well as on
the nature of the particle-liquid interface.
There are four related electrokinetic phenomena which are gen-
erally defined as follows: electrophoresis—the movement of a
charged surface (i.e., suspended particle) relative to a stationary liq-
uid induced by an applied electrical field, sedimentation potential—
the electric field which is crested when charged particles move
relative to a stationary liquid, electroosmosis—the movement of a
liquid relative to a stationary charged surface (i.e., capillary wall),
and streaming potential—the electric field which is created when
liquid is made to flow relative to a stationary charged surface. The
effects summarized by Eq. (20-23) form the basis of these electroki-
For many particles, the diffuse-charge layer can be characterized
adequately by the value of the zeta potential. For a spherical particle of
radius r0 which is large compared with the thickness of the diffuse-
charge layer, an electric field uniform at a distance from the particle
will produce a tangential electric field which varies with position on the
particle. Laplace’s equation [Eq. (20-19)] governs the distribution of
potential outside the diffuse-charge layer; also, the Navier-Stokes
equation for a creeping-flow regime can be applied to the velocity dis-
tribution. On account of the thinness of the diffuse-charge layer, Eq.
(20-23) can be used as a local boundary condition, accounting for the
effect of this charge in leading to movement of the particle relative to
the solution. The result of this computation gives the velocity of the
v = εζ ∇E/µ (20-24)
and it may be convenient to tabulate the mobility of the particle
U = v/∇E = εζ/µ (20-25)
rather than its zeta potential. Note that this mobility gives the velocity
of the particle for unit electric field rather than for unit force on the
particle. Related equations can be developed for the velocity of elec-
troosmotic flow. The subsections presented below (“Electrophoresis,”
“Electrofiltration,” and “Cross-Flow–Electrofiltration”) represent both
established and emerging commercial applications of electrokinetic
GENERAL REFERENCE: Wankat, Rate-Controlled Separations, Elsevier, Lon-
Electrophoretic Mobility Macromolecules move at speeds
measured in tenths of micrometers per second in a field (gradient) of
1 V/cm. Larger particles such as bubbles or bacteria move up to 10
times as fast because U is usually higher. To achieve useful separa-
tions, therefore, voltage gradients of 10 to 100 V/cm are required.
High voltage gradients are achieved only at the expense of power dis-
sipation within the fluid, and the resulting heat tends to cause unde-
sirable convection currents.
Mobility is affected by the dielectric constant and viscosity of the
suspending fluid, as indicated in Eq. (20-25). The ionic strength of the
fluid has a strong effect on the thickness of the double layer and hence
on ζ. As a rule, mobility varies inversely as the square root of ionic
strength [Overbeek, Adv. Colloid Sci., 3, 97 (1950)].
Modes of Operation There is a close analogy between sedimen-
tation of particles or macromolecules in a gravitational field and their
electrophoretic movement in an electric field. Both types of separa-
tion have proved valuable not only for analysis of colloids but also for
preparative work, at least in the laboratory. Electrophoresis is applica-
ble also for separating mixtures of simple cations or anions in certain
cases in which other separating methods are ineffectual.
Electrodecantation or electroconvection is one of several opera-
tions in which one mobile component (or several) is to be separated
out from less mobile or immobile ones. The mixture is introduced
between two vertical semipermeable membranes; for separating
cations, anion membranes are used, and vice versa. When an electric
field is applied, the charged component migrates to one or another of
the membranes; but since it cannot penetrate the membrane, it accu-
mulates at the surface to form a dense concentrated layer of particles
which will sink toward the bottom of the apparatus. Near the top of
the apparatus immobile components will be relatively pure. Murphy
[J. Electrochem. Soc., 97(11), 405 (1950)] has used silver-silver chlo-
ride electrodes in place of membranes. Frilette [J. Phys. Chem., 61,
168 (1957)], using anion membranes, partially separated H+
, and K+
Countercurrent electrophoresis can be used to split a mixture of
mobile species into two fractions by the electrical analog of elutria-
tion. In such countercurrent electrophoresis, sometimes termed an
ion still, a flow of the suspending fluid is maintained parallel to the
direction of the voltage gradient. Species which do not migrate fast
enough in the applied electric field will be physically swept out of the
apparatus. An apparatus based mainly on this principle but using also
natural convection currents has been developed (Bier, Electrophore-
sis, vol. II, Academic, New York, 1967).
Membrane electrophoresis which is based upon differences in ion
mobility, has been studied by Glueckauf and Kitt [J. Appl. Chem., 6,
511 (1956)]. Partial exclusion of coions by membranes results in large
differences in coion mobilities. Superposing a cation and an anion
membrane gives high transference numbers (about 0.5) for both
cations and anions while retaining the selectivity of mobilities. Large
voltages are required, and flow rates are low.
In continuous-flow zone electrophoresis the “solute” mixture to be
separated is injected continuously as a narrow source within a body of
carrier fluid flowing between two electrodes. As the “solute” mixture
passes through the transverse field, individual components migrate
sideways to produce zones which can then be taken off separately
downstream as purified fractions.
Resolution depends upon differences in mobilities of the species.
Background electrolyte of low ionic strength is advantageous, not only
to increase electrophoretic (solute) mobilities, but also to achieve low
electrical conductivity and thereby to reduce the thermal-convection
current for any given field [Finn, in Schoen (ed.), New Chemical
Engineering Separation Techniques, Interscience, New York, 1962].
The need to limit the maximum temperature rise has resulted in
two main types of apparatus, illustrated in Fig. 20-25. The first con-
sists of multicomponent ribbon separation units—apparatus capable
of separating small quantities of mixtures which may contain few or
many species. In general, such units operate with high voltages, low
20-20 ALTERNATIVE SEPARATION PROCESSES
currents, a large transverse dimension, and a narrow thickness be-
tween cooling faces. Numerous units developed for analytic chem-
istry, generally with filter-paper curtains but sometimes with granular
“anticonvectant” packing, are of this type. The second type consists of
block separation units—apparatus designed to separate larger quanti-
ties of a mixture into two (or at most three) species or fractions. Such
units generally use low to moderate voltages and high currents, with
cooling by circulation of cold electrolyte through the electrode com-
partments. Scale-up can readily be accomplished by extending the
thickness dimension w.
Both types of units have generally been operated in trace mode;
that is, “background” or “elutant” electrolyte is fed to the unit along
with the mixture to be separated. A desirable and possible means of
operation for preparative applications is in bulk mode, in which one
separated component follows the other without background elec-
trolyte being present, except that other ions may be required to
bracket the separated zones. Overlap regions between components
should be recycled, and pure components collected as products.
For block units, the need to stabilize flow has given rise to a num-
ber of distinct techniques.
Free flow. Dobry and Finn [Chem. Eng. Prog., 54, 59 (1958)]
used upward flow, stabilized by adding methyl cellulose, polyvinyl
alcohol, or dextran to the background solution. Upward flow was also
used in the electrode compartments, with cooling efficiency sufficient
to keep the main solution within 1°C of entering temperature.
Density gradients to stabilize flow have been employed by Philpot
[Trans. Faraday Soc., 36, 38 (1940)] and Mel [ J. Phys. Chem., 31,
559 (1959)]. Mel’s Staflo apparatus [ J. Phys. Chem., 31, 559 (1959)]
has liquid flow in the horizontal direction, with layers of increasing
density downward produced by sucrose concentrations increasing to
7.5 percent. The solute mixture to be separated is introduced in one
such layer. Operation at low electrolyte concentrations, low voltage
gradients, and low flow rates presents no cooling problem.
Packed beds. A packed cylindrical electrochromatograph 9 in
(23 cm) in diameter and 48 in (1.2 m) high, with operating voltages in
the 25- to 100-V range, has been developed by Hybarger, Vermeulen,
and coworkers [Ind. Eng. Chem. Process Des. Dev., 10, 91 (1971)].
The annular bed is separated from inner and outer electrodes by
porous ceramic diaphragms. The unit is cooled by rapid circulation of
cooled electrolyte between the diaphragms and the electrodes.
An interesting modification of zone electrophoresis resolves mix-
tures of ampholytes on the basis of differing isoelectric points
rather than differing mobilities. Such isoelectric spectra develop
when a pH gradient is established parallel to the electric field. Each
species then migrates until it arrives at the region of pH where it
possesses no net surface charge. A strong focusing effect is thereby
achieved [Kolin, in Glick (ed.), Methods of Biochemical Analysis,
vol. VI, Interscience, New York, 1958].
GENERAL REFERENCE: P. Krishnaswamy and P. Klinkowski, “Electrokinetics
and Electrofiltration,” in Advances in Solid-Liquid Separation, H. S. Murali-
dhara (ed.), Battelle Press, Columbus, OH, 1986.
Process Concept The application of a direct electric field of
appropriate polarity when filtering should cause a net charged-
particle migration relative to the filter medium (electrophoresis). The
same direct electric field can also be used to cause a net fluid flow
relative to the pores in a fixed filter cake or filter medium (electro-
osmosis). The exploitation of one or both of these phenomena form
the basis of conventional electrofiltration.
In conventional filtration, often the object is to form a high-solids-
content filter cake. At a single-filter surface, a uniform electric field
can be exploited in one of two ways. The first method of exploitation
occurs when the electric field is of a polarity such that the charged-
particle migration occurs toward the filter medium. In this case, the
application of the electric field increases the velocity of the solid par-
ticles toward the filter surface (electrosedimentation), thereby hasten-
ing the clarification of the feed suspension and, at the same time,
increasing the compaction of the filter cake collected on the filter sur-
face. In this first case, electroosmotic flow occurs in a direction away
from the filter media. The magnitude of the pressure-driven fluid flow
toward the filter surface far exceeds the magnitude of the electroos-
motic flow away from the surface so that the electroosmotic flow
results in only a minor reduction of the rate of production of filtrate.
The primary benefits of the applied electric field in this case are
increased compaction, and hence increased dewatering, of the filter
cake and an increased rate of sedimentation or movement of the par-
ticles in bulk suspension toward the filter surface.
The second method of exploitation occurs when the electric field is of
a polarity such that the charged-particle migration occurs away from the
filter medium. The contribution to the net-particle velocity of the elec-
trophoretically induced flow away from the filter medium is generally
orders of magnitude less than the contribution to the net-particle veloc-
ity of the flow induced by drag due to the pressure-induced flow of the
bulk liquid toward the filter media. (In conventional or cake filtration,
the velocity of liquid in dead-end flow toward the filter is almost always
sufficient to overcome any electrophoretic migration of particles away
from the filter media so that the prevention of the formation of filter
cake is not an option. This will not necessarily be the case for cross-flow
electrofiltration.) The primary enhancement to filtration caused by the
application of an electrical field in this manner is the increase in the fil-
trate flux due to electroosmotic flow through the filter cake. This elec-
troosmotic flow is especially beneficial during the latter stages of
filtration when the final filter-cake thickness has been achieved. At this
stage, electroosmosis can be exploited to draw filtrate out from the pore
structure of the filter cake. This type of drying of the filter cake is some-
times called electroosmotic dewatering.
Commercial Applications Krishnaswamy and Klinkowski, op. cit.,
describe the Dorr-Oliver EAVF®. The EAVF® combines vacuum filtra-
tion with electrophoresis and electroosmosis and has been described as a
series of parallel platelike electrode assemblies suspended in a tank con-
taining the slurry to be separated. When using the EAVF®, solids are
collected at both electrodes, one collecting a compacted cake simply by
electrophoretic attraction and the second collecting a compacted cake
though vacuum filtration coupled with electroosmotic dewatering. Upon
the completion of a collection cycle, the entire electrode assembly is
withdrawn from the slurry bath and the cake is removed. The EAVF® is
quoted as being best suited for the dewatering of ultrafine slurries (par-
ticle sizes typically less than 10 µm).
GENERAL REFERENCES: Henry, Lawler, and Kuo, Am. Inst. Chem. Eng. J.,
23(6), 851 (1977). Kuo, Ph.D. dissertation, West Virginia University, 1978.
Process Concept The application of a direct electric field of
appropriate polarity when filtering should cause a net charged-particle
ALTERNATIVE SOLID/LIQUID SEPARATIONS 20-21
FIG. 20-25 Types of arrangement for zone electrophoresis or electrochro-
matography. (a) Ribbon unit, with d > w; cooling at side faces. (b) Block unit,
with w > d; cooling at electrodes.
migration away from the filter medium. This electrophoretic migration
will prevent filter-cake formation and the subsequent reduction of fil-
ter performance. An additional benefit derived from the imposed elec-
tric field is an electroosmotic flux. The presence of this flux in the
membrane and in any particulate accumulation may further enhance
the filtration rate.
Cross-flow–electrofiltration (CF-EF) is the multifunctional separa-
tion process which combines the electrophoretic migration present in
electrofiltration with the particle diffusion and radial-migration forces
present in cross-flow filtration (CFF) (microfiltration includes cross-
flow filtration as one mode of operation in “Membrane Separation
Processes” which appears later in this section) in order to reduce fur-
ther the formation of filter cake. Cross-flow–electrofiltration can even
eliminate the formation of filter cake entirely. This process should find
application in the filtration of suspensions when there are charged
particles as well as a relatively low conductivity in the continuous
phase. Low conductivity in the continuous phase is necessary in order
to minimize the amount of electrical power necessary to sustain the
electric field. Low-ionic-strength aqueous media and nonaqueous
suspending media fulfill this requirement.
Cross-flow–electrofiltration has been investigated for both aqueous
and nonaqueous suspending media by using both rectangular-
and tubular-channel processing configurations (Fig. 20-26). Henry,
Lawler, and Kuo (op. cit.), using a rectangular-channel system with a
0.6-µ-pore-size polycarbonate Nuclepore filtration membrane, inves-
tigated CF-EF for 2.5-µm kaolin-water and 0.5- to 2-µm oil-in-water
emulsion systems. Kuo (op. cit.), using similar equipment, studied
5-µm kaolin-water, ∼100-µm Cr2O3-water, and ∼6-µm Al2O3-methanol
and/or -butanol systems. For both studies electrical fields of 0 to
60 V/cm were used for aqueous systems, and to 5000 V/cm were used
for nonaqueous systems. The studies covered a wide range of process-
ing variables in order to gain a better understanding of CF-EF funda-
mentals. Lee, Gidaspow, and Wasan [Ind. Eng. Chem. Fundam.,
19(2), 166 (1980)] studied CF-EF by using a porous stainless-steel
tube (pore size = 5 µm) as the filtration medium. A platinum wire run-
ning down the center of the tube acted as one electrode, while the
porous steel tube itself acted as the other electrode. Nonaqueous sus-
pensions of 0.3- to 2-µm Al2O3-tetralin and a coal-derived liquid
diluted with xylene and tetralin were studied. By operating with
applied electric fields (1000 to 10,000 V/cm) above the critical voltage,
clear particle-free filtrates were produced. It should be noted that the
pore size of the stainless-steel filter medium (5 µm) was greater than
the particle size of the suspended Al2O3 solids (0.3 to 2 µm). Cross-
flow–electrofiltration has also been applied to biological systems.
Brors, Kroner, and Deckwer [ECB6: Proc. 6th Eur. Cong. Biotech.,
511 (1994)] separated malate dehydrogenase from the cellular debris
of Baker’s yeast using CF-EF. A two- to fivefold increase in the spe-
cific enzyme transport rate was reported when electric field strengths
of 20 to 40 V/cm were used.
Theory Cross-flow–electrofiltration can theoretically be treated
as if it were cross-flow filtration with superimposed electrical effects.
These electrical effects include electroosmosis in the filter medium
and cake and electrophoresis of the particles in the slurry. The addi-
tion of the applied electric field can, however, result in some qualita-
tive differences in permeate-flux-parameter dependences.
The membrane resistance for CF-EF can be defined by specifying
two permeate fluxes as
Jom = ∆P/Rom (20-26)
Jm = ∆P/Rm (20-27)
where Jom is the flux through the membrane in the absence of an electric
field and any other resistance, m/s; Jm is the same flux in the presence of
an electric field; and Rom is the membrane resistance in the absence of an
electric field, (N⋅s)/m3
. When electroosmotic effects do occur,
Jm = Jom + Km E (20-28)
where Km is the electroosmotic coefficient of the membrane, m2
and E is the applied-electric-field strength, V/m. Equations (20-26),
(20-27), and (20-28) can be combined and rearranged to give Eq.
(20-29), the membrane resistance in the presence of an electric field.
Rm = (20-29)
Similarly, cake resistance can be represented as
Rc = (20-30)
where Joc is the flux through the cake in the absence of an electric field
or any other resistance, Roc is the cake resistance in the absence of an
electric field, and Kc is the electroosmotic coefficient of the cake. The
cake resistance is not a constant but is dependent upon the cake thick-
ness, which is in turn a function of the transmembrane pressure drop
and electrical-field strength.
Particulate systems require the addition of the term µe E in order to
account for the electrophoretic migration of the particle. The constant
µe is the electrophoretic mobility of the particle, m2
/(V⋅s). For the case
of the CF-EF, the film resistance Rf can be represented as
Rf = (20-31)
The resistances, when incorporated into equations descriptive of
cross-flow filtration, yield the general expression for the permeate flux
for particulate suspensions in cross-flow–electrofiltration systems.
There are three distinct regimes of operation in CF-EF. These
regimes (Fig. 20-27) are defined by the magnitude of the applied elec-
tric field with respect to the critical voltage Ec. The critical voltage is
defined as the voltage at which the net particle migration velocity
toward the filtration medium is zero. At the critical voltage, there is a
balance between the electrical-migration and radial-migration veloci-
ties away from the filter and the velocity at which the particles are swept
toward the filter by bulk flow. There is no diffusive transport at E = Ec
(Fig. 20-27b) because there is no gradient in the particle concentration
normal to the filter surface. At field strengths below the critical voltage
(Fig. 20-27a), all migration velocities occur in the same direction as in
the cross-flow-filtration systems discussed earlier. At values of applied
voltage above the critical voltage (Fig. 20-27c) qualitative differences
k ln ᎏ
ᎏ+ Ur + µeE
1 + ᎏ
1 + ᎏ
20-22 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-26 Alternative electrode configurations for cross-flow–electrofiltration.
E > Ec, increases in permeate flux rate are due only to electroosmosis in
the filtration medium.
One potential difficulty with CF-EF is the electrodeposition of
the particles at the electrode away from the filtration medium. This
phenomenon, if allowed to persist, will result in performance
decay of CF-EF with respect to maintenance of the electric field. Sev-
eral approaches such as momentary reverses in polarity, protection of
the electrode with a porous membrane or filter medium, and/or uti-
lization of a high fluid shear rate can minimize electrodeposition.
GENERAL REFERENCES: Gascoyne and Vykoukal, Electrophoresis, 23,
1973–1983 (2002). Pohl, in Moore (ed.), Electrostatics and Its Applications,
ALTERNATIVE SOLID/LIQUID SEPARATIONS 20-23
are observed. In this case, the electrophoretic-migration velocity away
from the filter medium is greater than the velocity caused by bulk flow
toward the filtration medium. Particles concentrate away from the filter
medium. This implies that particle concentration is lowest next to the
filter medium (in actuality, a clear boundary layer has been observed).
The influence of fluid shear still improves the transfer of particles down
the concentration gradient, but in this case it is toward the filtration
medium. When the particles are small and diffusive transport domi-
nates radial migration, increasing the circulation velocity will decrease
the permeate flux rate in this regime. When the particles are large and
radial migration dominates, the increase in circulation velocity will still
improve the filtration rate. These effects are illustrated qualitatively in
Fig. 20-28a. The solid lines represent systems in which the particle dif-
fusive effect dominates the radial-migration effect, while the dashed
lines represent the inverse. Figure 20-28b illustrates the increase in
filtration rate with increasing electric field strength. For field strengths
FIG. 20-27 Regimes of operation of cross-flow–electrofiltration: (a) voltage
less than critical, (b) voltage equal to the critical voltage, (c) voltage greater than
FIG. 20-28 Qualitative effects of Reynolds number and applied-electric-field
migration dominates); solid lines, small particles (particle diffusion dominates).
If, for example, the particle is more polarizable than the fluid, then the
net force is such as to impel the particle to regions of greater field
strength. Note that this statement implies that the effect is indepen-
dent of the absolute sign of the field direction. This is found to be the
case. Even rapidly alternating (ac) fields can be used to provide unidi-
rectional motion of the suspended particles.
Formal Theory A small neutral particle at equilibrium in a static
electric field experiences a net force due to DEP that can be written
as F = (p ⋅ ١)E, where p is the dipole moment vector and E is the
external electric field. If the particle is a simple dielectric and is
isotropically, linearly, and homogeneously polarizable, then the dipole
moment can be written as p = ␣vE, where ␣ is the (scalar) polariz-
ability, v is the volume of the particle, and E is the external field. The
force can then be written as:
F = ␣v(E ⋅ ١)E = a␣v١|E|2
This force equation can now be used to find the force in model sys-
tems such as that of an ideal dielectric sphere (relative dielectric con-
stant K2 ) in an ideal perfectly insulating dielectric fluid (relative
dielectric constant K1). The force can now be written as
F = 2πa3
ε0 K1 ١|E|2
(ideal dielectric sphere in ideal fluid).
Heuristic Explanation As we can see from Fig. 20-30, the DEP
response of real (as opposed to perfect insulator) particles with fre-
quency can be rather complicated. We use a simple illustration to
account for such a response. The force is proportional to the differ-
ence between the dielectric permittivities of the particle and the
surrounding medium. Since a part of the polarization in real systems
is thermally activated, there is a delayed response which shows as a
phase lag between D, the dielectric displacement, and E, the electric-
field intensity. To take this into account we may replace the simple
(absolute) dielectric constant ε by the complex (absolute) dielectric
constant ˆε = ε′ − iε″ = ε′ − iσ/w, where ω is the angular frequency of
the applied field. For treating spherical objects, for example, the
F ∝ → Re Ά · (20-34)
can be made, where ˆε° is the complex conjugate of ˆε.
With this force expression for real dielectrics, we can now explain
the complicated DEP response with the help of Fig. 20-30.
ˆε°1(ˆε2 − ˆε1)
ˆε2 + 2 ˆε1
ε1(ε2 − ε1)
ε2 + 2ε1
K2 − K1
K2 + 2K1
20-24 ALTERNATIVE SEPARATION PROCESSES
Wiley, New York, 1973, chap. 14 and chap. 15 (with Crane). Pohl, in Catsim-
poolas (ed.), Methods of Cell Separation, vol. I, Plenum Press, New York, 1977,
chap. 3. Pohl, Dielectrophoresis: The Behavior of Matter in Nonuniform Electric
Fields, Cambridge, New York, 1978.
Introduction Dielectrophoresis (DEP) is defined as the motion
of neutral, polarizable matter produced by a nonuniform electric (ac
or dc) field. DEP should be distinguished from electrophoresis, which
is the motion of charged particles in a uniform electric field (Fig.
The DEP of numerous particle types has been studied, and many
applications have been developed. Particles studied have included
aerosols, glass, minerals, polymer molecules, living cells, and cell
organelles. Applications developed include filtration, orientation,
sorting or separation, characterization, and levitation and materials
handling. Effects of DEP are easily exhibited, especially by large par-
ticles, and can be applied in many useful and desirable ways. DEP
effects can, however, be observed on particles ranging in size even
down to the molecular level in special cases. Since thermal effects
tend to disrupt DEP with molecular-sized particles, they can be con-
trolled only under special conditions such as in molecular beams.
Principle The principle of particle and cell separation, control, or
characterization by the action of DEP lies in the fact that a net force
can arise upon even neutral particles situated in a nonuniform electric
field. The force can be thought of as rising from the imaginary two-
step process of (1) induction or alignment of an electric dipole in a
particle placed in an electric field followed by (2) unequal forces on
the ends of that dipole. This arises from the fact that the force of an
electric field upon a charge is equal to the amount of the charge and
to the local field strength at that charge. Since the two (equal) charges
of the (induced or oriented) dipole of the particle lie in unequal field
strengths of the diverging field, a net force arises. If the particle is sus-
pended in a fluid, then the polarizability of that medium enters, too.
FIG. 20-30 A heuristic explanation of the dielectrophoretic-collection-rate
(DCR)-frequency spectrum. The curves for the absolute values of the complex
permittivities of the fluid medium and of the suspended particles are shown
lying nearly, but not entirely, coincident over the frequency range of the applied
electric field. When the permittivity (dielectric constant) of the particles
exceeds that of the suspending medium, the collection, or “positive dielec-
trophoresis,” occurs. In the frequency ranges in which the permittivity of the
particles is less than that of the suspending medium no collection at the regions
of higher field intensity occurs. Instead there is “negative dielectrophoresis,”
i.e., movement of the particles into regions of lower field intensity.
FIG. 20-29 Comparison of behaviors of neutral-charged bodies in an alternat-
ing nonuniform electric field. (a) Positively charged body moves toward nega-
tive electrode. Neutral body is polarized, then is attracted toward point where
field is strongest. Since the two charge regions on the neutral body are equal in
amount of charge but the force is proportional to the local field, a net force
toward the region of more intense field results. (b) Positively charged body
moves toward the negative electrode. Again, the neutral body is polarized, but it
does not reverse direction although the field is reversed. It still moves toward
the region of highest field intensity.
A particle, such as a living cell, can be imagined as having a number
of different frequency-dependent polarization mechanisms contribut-
ing to the total effective polarization of the particle |ˆε2|. The heavy
curve in Fig. 20-30 shows that the various mechanisms in the particle
drop out stepwise as the frequency increases. The light curve in Fig.
20-30 shows the polarization for a simple homogeneous liquid that
forms the surrounding medium. This curve is a smooth function
which becomes constant at high frequency. As the curves cross each
other (and hence |ε2| = |ε1|), various responses occur. The particle can
thus be attracted to the strongest field region, be repelled from that
region, or experience no force depending on the frequency.
Limitations It is desirable to have an estimate for the smallest
particle size that can be effectively influenced by DEP. To do this, we
consider the force on a particle due to DEP and also due to the
osmotic pressure. This latter diffusional force will randomize the par-
ticles and tend to destroy the control by DEP. Figure 20-31 shows a
plot of these two forces, calculated for practical and representative
conditions, as a function of particle radius. As we can see, the smallest
particles that can be effectively handled by DEP appear to be in range
of 0.01 to 0.1 µm (100 to 1000 Å).
Another limitation to be considered is the volume that the DEP
force can affect. This factor can be controlled by the design of elec-
trodes. As an example, consider electrodes of cylindrical geometry. A
practical example of this would be a cylinder with a wire running down
the middle to provide the two electrodes. The field in such a system is
proportional to 1/r. The DEP force is then FDEP ∝ ∇|E2
| ∝ 1/r3
, so that
any differences in particle polarization might well be masked merely
by positional differences in the force. At the outer cylinder the DEP
force may even be too small to affect the particles appreciably. The
most desirable electrode shape is one in which the force is indepen-
dent of position within the nonuniform field. This “isomotive” elec-
trode system is shown in Fig. 20-32.
Applications of Dielectrophoresis Over the past 20 years the
use of DEP has grown rapidly to a point at which it is in use for
biological [Hughes, Electrophoresis, 23, 2569–2582 (2002)], colloidal,
and mineral materials studies and handling. The effects of nonuni-
form electric fields are used for handling particulate matter far more
often than is usually recognized. This includes the removal of particu-
late matter by “electrofiltration,” the sorting of mixtures, or its con-
verse, the act of mixing, as well as the coalescence of suspensions. In
addition to these effects involving the translational motions of parti-
cles, some systems apply the orientational or torsional forces available
in nonuniform fields. One well-known example of the latter is the
placing of “tip-up” grit on emery papers commercially. Xerography
and many other imaging processes are examples of multibillion-dollar
industries which depend upon DEP for their success.
A clear distinction between electrophoresis (field action on an
object carrying excess free charges) and dielectrophoresis (field gra-
dient action on neutral objects) must be borne in mind at all times.
A dielectrofilter [Lin and Benguigui, Sep. Purif. Methods, 10(1), 53
(1981); Sisson et al., Sep. Sci. Technol., 30(7–9), 1421 (1995)] is a
device which uses the action of an electric field to aid the filtration and
removal of particulates from fluid media. A dielectrofilter can have a
very obvious advantage over a mechanical filter in that it can remove
particles which are much smaller than the flow channels in the filter.
In contrast, the ideal mechanical filter must have all its passages
smaller than the particles to be removed. The resultant flow resistance
can be use-restrictive and energy-consuming unless a phenomenon
such as dielectrofiltration is used.
Dielectrofiltration can (and often does) employ both electrophoresis
and dielectrophoresis in its application. The precise physical process
which dominates depends on a number of physical parameters of the
system. Factors such as field intensity and frequency and the electrical
conductivity and dielectric constants of the materials present deter-
mine this. Although these factors need constant attention for optimum
operation of the dielectrofilter, this additional complication is often
more than compensated for by the advantages of dielectrofiltration
such as greater throughput and lesser sensitivity to viscosity problems,
etc. To operate the dielectrofilter in the dominantly electrophoretic
mode requires that excess free charges of one sign or the other reside
on the particulate matter. The necessary charges can be those naturally
present, as upon a charged sol; or they may need to be artificially
implanted such as by passing the particles through a corona discharge.
ALTERNATIVE SOLID/LIQUID SEPARATIONS 20-25
FIG. 20-31 Comparison of the dielectrophoretic (Fd) and osmotic (Fos) forces
as functions of the particle size.
FIG. 20-32 A practical isomotive field geometry, showing r60, the critical
radius characterizing the isomotive electrodes. Electrode 3 is at ground poten-
tial, while electrodes 1 and 2 are at V1 = V+ and V2 = V− = −V+ respectively. The
inner faces of electrodes 1 and 2 follow r = r0 [sin (3θ/2)]−2/3
, while electrode 3
forms an angle of 120° about the midline.
Dielectrofiltration by the corona-charging, electrophoresis-dominated
Cottrell technique is now widely used.
To operate the dielectrofilter (dominantly dielectrophoretic mode),
on the other hand, one must avoid the presence of free charge on the
particles. If the particles can become charged during the operation, a
cycle of alternate charging and discharging in which the particles dash
to and from the electrodes can occur. This is most likely to occur if sta-
tic or very low frequency fields are used. For this reason, corona and
like effects may be troublesome and need often to be minimized. To be
sure, the DEP force is proportional to the field applied [actually to ∇
], but fields which are too intense can produce such troublesome
charge injection. A compromise for optimal operation is necessary
between having ∇(E)2
so low that DEP forces are insufficient for
dependable operation, on the one hand, and having E so high that
troublesome discharges (e.g., coronalike) interfere with dependable
operation of the dielectrofilter. In insulative media such as air or hydro-
carbon liquids, for example, one might prefer to operate with fields in
the range of, say, 10 to 10,000 V/cm. In more conductive media such as
water, acetone, or alcohol, for example, one would usually prefer rather
lower fields in the range of 0.01 to 100 V/cm. The higher field ranges
cited might become unsuitable if conductive sharp asperities are
Another factor of importance in dielectrofiltration is the need to
have the DEP effect firmly operative upon all portions of the fluid
passing through. Oversight of this factor is a most common cause of
incomplete dielectrofiltration. Good dielectrofilter design will
emphasize this crucial point. To put this numerically, let us consider
the essential field factor for DEP force, namely ∇(E0)2
. Near sharp
points, e.g., E, the electric field varies with the radial distance r as E ∝
; hence our DEP force factor will vary as ∇(E)2
. In the neigh-
borhood of sharp “line” sources such as at the edge of electrode plates,
E ∝ r−1
, hence, ∇(E)2
. If, for instance, the distance is varied by a
factor of 4 from the effective field source in these cases, the DEP
force can be expected to weaken by a factor of 1024 or 64 respectively
for the point source and the line source. The matter is even more
keenly at issue when field-warping dielectrics (defined later) are used
to effect maximal filtration. In this case the field-warping material is
made to produce dipole fields as induced by the applied electric field.
If we ask how the crucial factor, ∇(E)2
, varies with distance away from
such a dipole, we find that since the field Ed about a dipole varies
approximately as r−3
, then ∇(E)2
can be expected to vary as r−7
. It then
becomes critically important that the particles to be removed from the
passing fluid do, indeed, pass very close to the surface of the field-
warping material, or it will not be effectively handled. Clearly, it would
be difficult to maintain successfully uniform dielectrofiltration treat-
ment of fluid passing through such wildly variant regions. The prob-
lems can be minimized by ensuring that all the elements of the passing
fluid go closely by such field sources in the dielectrofilter. In practice
this is done by constructing the dielectrofilter from an assembly of
highly comminuted electrodes or else by a set of relatively simple and
widely spaced metallic electrodes between which is set an assembly of
more or less finely divided solid dielectric material having a complex
permittivity different from that of the fluid to be treated. The solid
dielectric (fibers, spheres, chunks) serves to produce field nonunifor-
mities or field warpings to which the particles to be filtered are to be
attracted. In treating fluids of low dielectric constant such as air or
hydrocarbon fluids, one sees field-warping materials such as sintered
ceramic balls, glass-wool matting, open-mesh polyurethane foam, alu-
mina, chunks, or BaTiO3 particles.
An example of a practical dielectrofilter which uses both of the fea-
tures described, namely, sharp electrodes and dielectric field-warping
filler materials, is that described in Fig. 20-33 [H. J. Hall and R. F.
Brown, Lubric. Eng., 22, 488 (1966)]. It is intended for use with
hydraulic fluids, fuel oils, lubricating oils, transformer oils, lubricants,
and various refinery streams. Performance data are cited in Fig.
20-34. It must be remarked that in the opinion of Hall and Brown the
action of the dielectrofilter was “electrostatic” and due to free charge
on the particles dispersed in the liquids. It is the present authors’
opinion, however, that both electrophoresis and dielectrophoresis are
operative here but that the dominant mechanism is that of DEP, in
which neutral particles are polarized and attracted to the regions of
highest field intensity.
20-26 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-33 Diagram illustrating the function of an electrostatic liquid cleaner.
FIG. 20-34 Performance data for a typical high-efficiency electrostatic liquid
ALTERNATIVE SOLID/LIQUID SEPARATIONS 20-27
A second commercial example of dielectric filtration is the
Gulftronic® separator [G. R. Fritsche, Oil & Gas J., 75, 73 (1977)] which
was commercialized in the late 1970s by Gulf Science and Technology
Company. Instead of using needle-point electrodes as shown in Fig. 20-33,
the Gulftronic® separator relied on the use of a bed of glass beads to pro-
duce the field nonuniformities required for dielectric filtration. Either ac
or dc electric fields could be used in this separator. The Gulftronic® sep-
arator has been used primarily to remove catalyst fines from FCC decant
oils and has been reported to exhibit removal efficiencies in excess of
80 percent for this fine-particle separation problem.
Another example of the commercial use of DEP is in polymer clarifi-
cation [A. N. Wennerberg, U.S. Patent 2,914,453, 1959; assignor to
Standard Oil Co. (Indiana)]. Here, either ac or dc potentials were used
while passing suspensions to be clarified through regions with an area-
to-electrode-area ratio of 10:1 or 100:1 and with fields in the order of
10 kV/cm. Field warping by the presence of various solid dielectrics was
observed to enhance filtration considerably, as expected for DEP. The
filtration of molten or dissolved polymers to free them of objectionable
quantities of catalyst residues, for example, was more effective if a solid
dielectric material such as Attapulgus clay, silica gel, fuller’s earth, alu-
mina, or bauxite was present in the region between the electrodes. The
effectiveness of percolation through such absorptive solids for removing
color bodies is remarkably enhanced by the presence of an applied field.
A given amount of clay is reported to remove from 4 to 10 times as
much color as would be removed in the absence of DEP. Similar results
are reported by Lin et al. [Lin, Yaniv, and Zimmels, Proc. XIIIth Int.
Miner. Process. Congr., Wroc⁄law, Poland, 83–105 (1979)].
The instances cited were examples of the use of DEP to filter liq-
uids. We now turn to the use of DEP to aid in dielectrofiltration of
gases. Fielding et al. observe that the effectiveness of high-quality
fiberglass air filters is dramatically improved by a factor of 10 or more
by incorporating DEP in the operation. Extremely little current or
power is required, and no detectable amounts of ozone or corona need
result. The DEP force, once it has gathered the particles, continues to
TABLE 20-14 Dielectrophoretic Augmentation of Filtration
of a Liquid Aerosol*
cm/s 2 kV 3.5 kV 5 kV 7 kV
0.3-µm-diameter dioctyl phthalate aerosol
3 8 19 95 330
6 3 13 39 120
9 3 11 28 100
15 2 6 13 42
20 2 5 9 27
28 2 4 6 14
39 2 3 4 9
50 1 2 3 6
1.0-µm-diameter dioctyl phthalate aerosol
3 30 110 300 1100
6 6 3 95 360
9 4 18 50 170
15 3 10 20 50
20 2 6 13 35
28 2 4 8 18
39 2 3 5 11
50 1 2 3 7
6 10 30 80
10 8 30 80
14 5 20 40
20 4 10 30 70
35 3 7 10 20
45 1 2 6
53 1 2 7 10
*Experimentally measured dielectrophoretic augmentation factor DAF as a
function of air speed and applied voltage for a glass-fiber filter (HP-100, Farr
Co.). Cf. Fielding, Thompson, Bogardus, and Clark, Dielectrophoretic Filtra-
tion of Solid and Liquid Aerosol Particulates, Prepr. 75-32.2, 68th ann. meet.,
Air Pollut. Control Assoc., Boston, June 1975.
act on the particles already sitting on the filter medium, thereby
improving adhesion and minimizing blowoff.
The degree by which the DEP increases the effectiveness of gas fil-
tration, or the dielectrophoretic augmentation factor (DAF), is defin-
able. It is the ratio of the volumes of aerosol-laden gas which can be
cleaned effectively by the filter with and without the voltage applied.
For example, the application of 11 kV/cm gave a DAF of 30 for
1.0-µm-diameter dioctyl phthalate particles in air, implying that the
penetration of the glass filter is reduced thirtyfold by the application of
a field of 1100 kV/m. Similar results were obtained by using “standard”
fly ash supplied by the Air Pollution Control Office of the U.S. Envi-
ronmental Protection Agency. The data obtained for several aerosols
tested are shown in Table 20-14 and in Fig. 20-35. The relation DAF =
/v is observed to hold approximately for each aerosol. Here, the
DEP augmentation factor DAF is observed to depend upon a constant
K, a characteristic of the material, upon the square of the applied volt-
age, and upon the inverse of the volume flow rate v through the filter.
It is worth noting that in the case of the air filter described DEP serves
as an augmenting rather than as an exclusive mechanism for the removal
of particulate material. It is a unique feature of the dielectrophoretic gas
filter that the DEP force is maximal when the particulates are at or on
the fiber surface. This causes the deposits to be strongly retained by this
particular filtration mechanism. It thus contrasts importantly with other
types of gas filter in which the filtration mechanism no longer acts after
the capture of the particle. In particular, in the case of the older electro-
static mechanisms involving only coulombic attraction, a simple charge
alternation on the particle, such as caused by normal conduction, often
evokes disruption of the filter operation because of particle repulsion
from the contacting electrode. On the other hand, ordinary mechanical
filtration depends upon the action of adventitious particle trapping or
upon van der Waals forces, etc., to hold the particles. The high efficiency
possible with electrofilters suggests their wider use.
FIG. 20-35 Efficiency of an electrofilter as a function of gas flow rate at 5 dif-
ferent voltages. Experimental materials: 1-µm aerosol of dioctyl phthalate; glass-
fiber filter. Symbols: ᭺, no voltage applied; ∆, 2 kV; ●, 3.5 kV; ᮀ, 5 kV; ▲, 7 kV.
(After Fielting et al., Dielectrophoretic Filtration of Solid and Liquid Aerosol
Particulates, Prepr. 75-32.2, 68th ann. meet., Air Pollut. Control Assoc., Boston,
fractionate mixtures of biological products. (See also Sec. 15.) In order
to demonstrate the versatility of particle distribution, he has cited the
example shown in Table 20-15. The feed mixture consisted of poly-
styrene particles, red blood cells, starch, and cellulose. Liquid-liquid
particle distribution has also been studied by using mineral-matter par-
ticles (average diameter = 5.5 µm) extracted from a coal liquid as the
solid in a xylene-water system [Prudich and Henry, Am. Inst. Chem.
Eng. J., 24(5), 788 (1978)]. By using surface-active agents in order to
enhance the water wettability of the solid particles, recoveries of better
than 95 percent of the particles to the water phase were observed. All
particles remained in the xylene when no surfactant was added.
Particle collection at a liquid-liquid interface is a particularly favor-
able separation process when applied to fine-particle systems. Advan-
tages of this type of processing include:
• Decreased liquid-liquid interfacial tension (when compared with a
gas-liquid system) results in higher liquid-liquid interfacial areas,
which favor solid-particle droplet collisions.
• Liquid-solid interactions due to long-range intermolecular forces
are much larger than are gas-solid interactions. This means that it is
easier to collect fine particles at a liquid-liquid interface than at a
• The increased momentum of liquid droplets (when compared with
gas) should favor solid-particle collection.
Fuerstenau [Lai and Fuerstenau, Trans. Am. Inst. Min. Metall. Pet.
Eng., 241, 549 (1968); Raghavan and Fuerstenau, Am. Inst. Chem.
Eng. Symp. Ser., 71(150), 59 (1975)] has studied this process with
respect to the removal of alumina particles (0.1 µm) and hematite par-
ticles (0.2 µm) from an aqueous solution by using isooctane. The use
of isooctane as the collecting phase for the hematite particles resulted
in an increase in particle recovery of about 50 percent over that mea-
sured when air was used as the collecting phase under the same con-
ditions. The effect of the wettability of the solid particles (as measured
by the three-phase contact angle) on the recovery of hematite in the
water-isooctane system is shown in Fig. 20-37. This behavior is typical
of particle collection. Particle collection at an oil-water interface has
also been studied with respect to particle removal from a coal liquid.
Particle removals averaging about 80 percent have been observed
when water is used as the collecting phase (Lau, master’s thesis, West
Virginia University, 1979). Surfactant addition was necessary in order
to control the wettability of the solids.
Particle bridging has been chiefly investigated with respect to
spherical agglomeration. Spherical agglomeration involves the col-
lecting or transferring of the fine particles from suspension in a liquid
phase into spherical aggregates held together by a second liquid
phase. The aggregates are then removed from the slurry by filtration
or settling. Like the other liquid-solid-liquid separation techniques,
the solid must be wet by the second liquid phase. The spherical
agglomeration process has resulted in the development of a pilot unit
called the Shell Pelletizing Separator [Zuiderweg and Van Lookeren
Campagne, Chem. Eng. (London), 220, CE223 (1968)].
The ability to determine in advance which of the separation
regimes is most advantageous for a given liquid-solid-liquid system
would be desirable. No set of criteria with which to make this deter-
mination presently exists. Work has been done with respect to the
20-28 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-36 Regimes of separation in a liquid-solid-liquid system. Phase 1 =
particle; phase 2 = liquid (dispersed); phase 3 = liquid (continuous).
TABLE 20-15 Separations of Particles between Two Phases
System Top phase Bottom phase
Polystyrene All others
Algae All others
Red cells All others
Cellulose particles Starch
Dextran; 200,000 MW
Dextran; 20,000 MW
SURFACE-BASED SOLID-LIQUID SEPARATIONS
INVOLVING A SECOND LIQUID PHASE
GENERAL REFERENCES: Fuerstenau, “Fine Particle Flotation,” in Somasun-
daran (ed.), Fine Particles Processing, vol. 1, American Institute of Mining,
Metallurgical, and Petroleum Engineers, New York, 1980. Henry, Prudich, and
Lau, Colloids Surf., 1, 335 (1980). Henry, Prudich, and Vaidyanathan, Sep.
Purif. Methods, 8(2), 31 (1979). Jacques, Hovarongkura, and Henry, Am. Inst.
Chem. Eng. J., 25(1), 160 (1979). Stratton-Crawley, “Oil Flotation: Two Liquid
Flotation Techniques,” in Somasundaran and Arbiter (eds.), Beneficiation of
Mineral Fines, American Institute of Mining, Metallurgical, and Petroleum
Engineers, New York, 1979.
Process Concept Three potential surface-based regimes of sep-
aration exist when a second, immiscible liquid phase is added to
another, solids-containing liquid in order to effect the removal of
solids. These regimes (Fig. 20-36) are:
1. Distribution of the solids into the bulk second liquid phase
2. Collection of the solids at the liquid-liquid interface
3. Bridging or clumping of the solids by the added fluid in order to
form an agglomerate followed by settling or filtration
These separation techniques should find particular application in
systems containing fine particles. The surface chemical differences
involved among these separation regimes are only a matter of degree;
i.e., all three regimes require the wetting of the solid by the second
liquid phase. The addition of a surface-active agent is sometimes
needed in order to achieve the required solids wettability. In spite of
this similarity, applied processing (equipment configuration, operat-
ing conditions, etc.) can vary widely. Collection at the interface would
normally be treated as a flotation process (see also “Adsorptive-
Bubble Separation Methods” in Sec. 20), distribution to the bulk liq-
uid as a liquid-liquid extraction analog, and particle bridging as a
settling (sedimentation) or filtration process.
Even though surface-property-based liquid-solid-liquid separation
techniques have yet to be widely used in significant industrial applica-
tions, several studies which demonstrate their effectiveness have
appeared in literature.
Albertsson (Partition of Cell Particles and Macromolecules, 3d ed.,
Wiley, New York, 1986) has extensively used particle distribution to
identification of system parameters which make these processes tech-
nically feasible. The results of these studies can be used to guide the
selection of the second liquid phase as well as to suggest approximate
operating conditions (dispersed-liquid droplet size, degree and type of
mixing, surface-active-chemical addition, etc.).
Theory Theoretical analyses of spherical particles suspended in a
planar liquid-liquid interface have appeared in literature for some
time, the most commonly presented forms being those of a free
energy and/or force balance made in the absence of all external body
forces. These analyses are generally used to define the boundary cri-
teria for the shift between the collection and distribution regimes, the
bridging regime not being considered. This type of analysis shows that
for a spherical particle possessing a three-phase contact angle
between 0 and 180°, as measured through the receiving or collecting
phase, collection at the interface is favored over residence in either
bulk phase. These equations are summarized, using a derivation of
Young’s equation, as
> 1 particles wet to phase 1 (20-35)
ᎏ < −1 particles wet to phase 2 (20-36)
ᎏΈ≤ 1 particle at interface (20-37)
where γij is the surface tension between phases i and j, N/m (dyn/cm);
s indicates the solid phase; and subscripts 1 and 2 indicate the two liq-
Several additional studies [Winitzer, Sep. Sci., 8(1), 45 (1973); ibid.,
8(6), 647 (1973); Maru, Wasan, and Kintner, Chem. Eng. Sci., 26,
1615 (1971); and Rapacchietta and Neumann, J. Colloid Interface
Sci., 59(3), 555 (1977)] which include body forces such as gravita-
tional acceleration and buoyancy have been made. A typical example
of a force balance describing such a system (Fig. 20-38) is summarized
in Eq. (20-38):
[(γs1 − γs2) cos δ + γ12 cos B]L = g[Vtotalρs − V1ρ1 − V2ρ2] (20-38)
where V1 is the volume of the particle in fluid phase 1, V2 is the volume
in fluid phase 2, L is the particle circumference at the interface between
the two liquid phases, ρi is the density of phase i, and g is the gravitational
constant. The left-hand side of the equation represents the surface
forces acting on the solid particle, while the right-hand side includes the
gravitational and buoyancy forces. This example illustrates the fact that
body forces can have a significant effect on system behavior. The solid-
particle size as well as the densities of the solid and both liquid phases are
introduced as important system parameters.
A study has also been performed for particle distribution for cases in
which the radii of curvature of the solid and the liquid-liquid interface
γs2 − γs1
are of the same order of magnitude [Jacques, Hovarongkura, and
Henry, Am. Inst. Chem. Eng. J., 25(1), 160 (1979)]. Differences
between the final and initial surface free energies are used to analyze
this system. Body forces are neglected. Results (Fig. 20-39) demon-
strate that n, the ratio of the particle radius to the liquid-liquid-interface
radius, is an important system parameter. Distribution of the particle
from one phase to the other is favored over continued residence in the
original phase when the free-energy difference is negative. For a solid
particle of a given size, these results show that as the second-phase
droplet size decreases, the contact angle required in order to effect dis-
tribution decreases (the required wettability of the solid by the second
phase increases). The case of particle collection at a curved liquid-liquid
interface has also been studied in a similar manner [Smith and Van de
Ven, Colloids Surf., 2, 387 (1981)]. This study shows that collection is
preferred over distribution for any n in systems without external body
forces when the contact angle lies between 0 and 180°.
While thermodynamic-stability studies can be valuable in evaluat-
ing the technical feasibility of a process, they are presently inadequate
in determining which separation regime will dominate a particular
liquid-solid-liquid system. These analyses ignore important process-
ing phenomena such as the mechanism of encounter of the dispersed-
phase liquid with the solid particles, the strength of particle attachment,
and the mixing-energy input necessary to effect the separation. No
models of good predictive value which take all these variables into
account have yet been offered. Until the effects of these and other sys-
tem variables can be adequately understood, quantified, and combined
into such a predictive model, no a priori method of performance pre-
diction will be possible.
ADSORPTIVE-BUBBLE SEPARATION METHODS
GENERAL REFERENCES: Lemlich (ed.), Adsorptive Bubble Separation Tech-
niques, Academic, New York, 1972. Carleson, “Adsorptive Bubble Separation
Processes” in Scamehorn and Harwell (eds.), Surfactant-Based Separation
Processes, Marcel Dekker, New York, 1989.
Principle The adsorptive-bubble separation methods, or adsub-
ble methods for short [Lemlich, Chem. Eng. 73(21), 7 (1966)], are
based on the selective adsorption or attachment of material on the
surfaces of gas bubbles passing through a solution or suspension. In
most of the methods, the bubbles rise to form a foam or froth which
carries the material off overhead. Thus the material (desirable or
undesirable) is removed from the liquid, and not vice versa as in, say,
ALTERNATIVE SOLID/LIQUID SEPARATIONS 20-29
FIG. 20-37 The variation of adsorption density, oil-droplet contact angle, and
oil-extraction recovery of hematite as a function of pH. To convert gram-moles
per square centimeter to pound-moles per square foot, multiply by 2.048. [From
Raghavan and Fuerstenau, Am. Inst. Chem. Eng. Symp. Ser., 71(150), 59
FIG. 20-38 Solid sphere suspended at the liquid-liquid interface. F1 and F2
are buoyancy forces; FS is gravity. [From Winitzer, Sep. Sci., 8(1), 45 (1973).]
FIG. 20-39 Normalized free-energy difference between distributed (II) and nondistributed (I) states of the solid particles versus three-phase contact
angle (collection at the interface is not considered). A negative free-energy difference implies that the distributed state is preferred over the nondistrib-
uted state. Note especially the significant effect of n, the ratio of the liquid droplet to solid-particle radius. [From Jacques, Hovarongkura, and Henry, Am.
Inst. Chem. Eng. J., 25(1), 160 (1979).]
filtration. Accordingly, the foaming methods appear to be particularly
(although not exclusively) suited to the removal of small amounts of
material from large volumes of liquid.
For any adsubble method, if the material to be removed (termed
the colligend) is not itself surface-active, a suitable surfactant
(termed the collector) may be added to unite with it and attach or
adsorb it to the bubble surface so that it may be removed (Sebba, Ion
Flotation, Elsevier, New York, 1962). The union between colligend
and collector may be by chelation or other complex formation. Alter-
natively, a charged colligend may be removed through its attraction
toward a collector of opposite charge.
Definitions and Classification Figure 20-40 outlines the most
widely accepted classification of the various adsubble methods
[Karger, Grieves, Lemlich, Rubin, and Sebba, Sep. Sci., 2, 401
(1967)]. It is based largely on actual usage of the terms by various
workers, and so the definitions include some unavoidable inconsisten-
cies and overlap.
Among the methods of foam separation, foam fractionation usu-
ally implies the removal of dissolved (or sometimes colloidal) material.
The overflowing foam, after collapse, is called the foamate. The solid
lines of Fig. 20-41 illustrate simple continuous foam fractionation.
(Batch operation would be represented by omitting the feed and bot-
On the other hand, flotation usually implies the removal of solid
particulate material. Most important under the latter category is ore
Also under the category of flotation are to be found macroflota-
tion, which is the removal of macroscopic particles; microflotation
(also called colloid flotation), which is the removal of microscopic
particles, particularly colloids or microorganisms [Dognon and
Dumontet, Comptes Rendus, 135, 884 (1941)]; molecular flotation,
which is the removal of surface-inactive molecules through the use of a
collector (surfactant) which yields an insoluble product; ion flotation,
which is the removal of surface-inactive ions via a collector which yields
an insoluble product, especially a removable scum [Sebba, Nature,
184, 1062 (1959)]; adsorbing colloid flotation, which is the removal
of dissolved material in piggyback fashion by adsorption on colloidal
particles; and precipitate flotation, in which a precipitate is removed
by a collector which is not the precipitating agent [Baarson and Ray,
“Precipitate Flotation,” in Wadsworth and Davis (eds.), Unit Processes
in Hydrometallurgy, Gordon and Breach, New York, 1964, p. 656].
The last definition has been narrowed to precipitate flotation of the
first kind, the second kind requiring no separate collector at all [Mahne
and Pinfold, J. Appl. Chem., 18, 52 (1968)].
A separation can sometimes be obtained even in the absence of any
foam (or any floated floc or other surrogate). In bubble fractiona-
tion this is achieved simply by lengthening the bubbled pool to form
a vertical column [Dorman and Lemlich, Nature, 207, 145 (1965)].
The ascending bubbles then deposit their adsorbed or attached mate-
rial at the top of the pool as they exit. This results in a concentration
gradient which can serve as a basis for separation. Bubble fractiona-
tion can operate either alone or as a booster section below a foam
20-30 ALTERNATIVE SEPARATION PROCESSES
fractionator, perhaps to raise the concentration up to the foaming
In solvent sublation an immiscible liquid is placed atop the main
liquid to trap the material deposited by the bubbles as they exit
(Sebba, Ion Flotation, Elsevier, New York, 1962). The upper liquid
should dissolve or at least wet the material. With appropriate selectiv-
ity, the separation so achieved can sometimes be much greater than
that with bubble fractionation alone.
The droplet analogs to the adsubble methods have been termed the
adsoplet methods (from adsorptive droplet separation methods)
[Lemlich, “Adsorptive Bubble Separation Methods,” Ind. Eng. Chem.,
60(10), 16 (1968)]. They are omitted from Fig. 20-40, since they involve
adsorption or attachment at liquid-liquid interfaces. Among them are
emulsion fractionation [Eldib, “Foam and Emulsion Fractionation,”
in Kobe and McKetta (eds.), Advances in Petroleum Chemistry and
Refining, vol. 7, Interscience, New York, 1963, p. 66], which is the ana-
log of foam fractionation; and droplet fractionation [Lemlich, loc.
cit.; and Strain, J. Phys. Chem., 57, 638 (1953)], which is the analog of
bubble fractionation. Similarly, the old beneficiation operation called
bulk oil flotation (Gaudin, Flotation, 2d ed., McGraw-Hill, New York,
1957) is the analog of modern ore flotation. By and large, the adsoplet
methods have not attracted the attention accorded to the adsubble
Of all the adsubble methods, foam fractionation is the one for
which chemical engineering theory is the most advanced. Fortunately,
some of this theory also applies to other adsubble methods.
Adsorption The separation achieved depends in part on the
selectivity of adsorption at the bubble surface. At equilibrium, the
adsorption of dissolved material follows the Gibbs equation (Gibbs,
Collected Works, Longmans Green, New York, 1928).
dγ = −RTΣΓi d ln ai (20-39)
Γi is the surface excess (Davies and Rideal, Interfacial Phenomena, 2d
ed., Academic, New York, 1963). For most purposes, it is sufficient to
view Γi as the concentration of adsorbed component i at the surface in
units of, say (g⋅mol)/cm2
. R is the gas constant, T is the absolute temper-
ature, γ is the surface tension, and ai is the activity of component i. The
minus sign shows that material which concentrates at the surface gener-
ally lowers the surface tension, and vice versa. This can sometimes be a
guide in determining preliminarily what materials can be separated.
When applied to a nonionic surfactant in pure water at concentra-
tions below the critical micelle concentration, Eq. (20-39) simplifies
into Eq. (20-40)
Γs = − (20-40)
C is the concentration in the bulk, and subscript s refers to the surfac-
tant. Under some conditions, Eq. (20-40) may apply to an ionic sur-
factant as well (Lemlich, loc. cit.).
The major surfactant in the foam may usually be considered to be
present at the bubble surfaces in the form of an adsorbed monolayer
with a substantially constant Γs, often of the order of 3 × 10−10
, for a molecular weight of several hundred. On the other hand,
trace materials follow the linear-adsorption isotherm Γi = KiCi if their
concentration is low enough. For a wider range of concentration a
Langmuir or other type of isotherm may be applicable (Davies and
Rideal, loc. cit.).
Factors Affecting Adsorption Ki for a colligend can be
adversely affected (reduced) through an insufficiency of collector. It
can also be reduced through an excess of collector, which competes
d ln Cs
ALTERNATIVE SOLID/LIQUID SEPARATIONS 20-31
FIG. 20-40 Classification for the adsorptive-bubble separation methods.
FIG. 20-41 Four alternative modes of continuous-flow operation with a foam-
fractionation column: (1) The simple mode is illustrated by the solid lines. (2)
Enriching operation employs the dashed reflux line. (3) In stripping operation,
the elevated dashed feed line to the foam replaces the solid feed line to the pool.
(4) For combined operation, reflux and elevated feed to the foam are both
for the available surface against the collector-colligend complex
[Schnepf, Gaden, Mirocznik, and Schonfeld, Chem. Eng. Prog.,
55(5), 42 (1959)].
Excess collector can also reduce the separation by forming micelles
in the bulk which adsorb some of the colligend, thus keeping it from
the surface. This effect of the micelles on Ki for the colligend is given
theoretically [Lemlich, “Principles of Foam Fractionation,” in Perry
(ed.), Progress in Separation and Purification, vol. 1, Interscience,
New York, 1968, chap. 1] by Eq. (20-41) [Lemlich (ed.), Adsorptive
Bubble Separation Techniques, Academic, New York, 1972] if Γs is
constant when Cs > Csc:
= + (20-41)
K1 is Ki just below the collector’s critical micelle concentration, Csc. K2
is Ki at some higher collector concentration, Cs. E is the relative effec-
tiveness, in adsorbing colligend, of surface collector versus micellar
collector. Generally, E > 1. Γs is the surface excess of collector. More
about each K is available [Lemlich, “Adsubble Methods,” in Li (ed.),
Recent Developments in Separation Science, vol. 1, CRC Press, Cleve-
land, 1972, pp. 113–127; Jashnani and Lemlich, Ind. Eng. Chem.
Process Des. Dev., 12, 312 (1973)].
The controlling effect of various ions can be expressed in terms
of thermodynamic equilibria [Karger and DeVivo, Sep. Sci., 3, 393
1968)]. Similarities with ion exchange have been noted. The selectiv-
ity of counterionic adsorption increases with ionic charge and de-
creases with hydration number [Jorne and Rubin, Sep. Sci., 4, 313
(1969); and Kato and Nakamori, J. Chem. Eng. Japan, 9, 378 (1976)].
By analogy with other separation processes, the relative distribution in
multicomponent systems can be analyzed in terms of a selectivity
coefficient αmn = ΓmCn /ΓnCm [Rubin and Jorne, Ind. Eng. Chem. Fun-
dam., 8, 474 (1969); J. Colloid Interface Sci., 33, 208 (1970)].
Operation in the Simple Mode If there is no concentration
gradient within the liquid pool and if there is no coalescence within
the rising foam, then the operation shown by the solid lines of Fig.
20-41 is truly in the simple mode, i.e., a single theoretical stage of sep-
aration. Equations (20-42) and (20-43) will then apply to the steady-
CQ = CW + (GSΓW /Q) (20-42)
CW = CF − (GSΓW /F) (20-43)
CF, CW, and CQ are the concentrations of the substance in question
(which may be a colligend or a surfactant) in the feed stream, bottoms
stream, and foamate (collapsed foam), respectively. G, F, and Q are
the volumetric flow rates of gas, feed, and foamate, respectively. ΓW is
the surface excess in equilibrium with CW. S is the surface-to-volume
ratio for a bubble. For a spherical bubble, S = 6/d, where d is the bub-
ble diameter. For variation in bubble sizes, d should be taken as
, where ni is the number of bubbles with diameter di in a
representative region of foam.
Finding G Either Eq. (20-42) or Eq. (20-43) can be used to find the
surface excess indirectly from experimental measurements. To assure a
close approach to operation as a single theoretical stage, coalescence in
the rising foam should be minimized by maintaining a proper gas rate
and a low foam height [Brunner and Lemlich, Ind. Eng. Chem. Fundam.
2, 297 (1963)]. These precautions apply particularly with Eq. (20-42).
For laboratory purposes it is sometimes convenient to recycle the
foamate directly to the pool in a manner analogous to an equilibrium
still. This eliminates the feed and bottoms streams and makes for a
more reliable approach to steady-state operation. However, this re-
cycling may not be advisable for colligend measurements in the pres-
ence of slowly dissociating collector micelles.
To avoid spurious effects in the laboratory, it is advisable to employ
a prehumidified chemically inert gas.
Bubble Sizes Subject to certain errors (de Vries, Foam Stability,
Rubber-Stichting, Delft, 1957), foam bubble diameters can be measured
photographically. Some of these errors can be minimized by taking pains
to generate bubbles of fairly uniform size, say, by using a bubbler with
identical orifices or by just using a bubbler with a single orifice (gas rate
permitting). Otherwise, a correction for planar statistical sampling bias in
Cs − Csc
the foam should be incorporated with actual diameters [de Vries, op. cit.]
or truncated diameters [Lemlich, Chem. Eng. Commun. 16, 153 (1982)].
Also, size segregation can reduce mean mural bubble diameter by
roughly half the standard deviation [Cheng and Lemlich, Ind. Eng.
Chem. Fundam. 22, 105 (1983)]. Bubble diameters can also be mea-
sured in the liquid pool, either photographically or indirectly via mea-
surement of the gas flow rate and stroboscopic determination of bubble
frequency [Leonard and Lemlich, Am. Inst. Chem. Eng. J., 11, 25
Bubble sizes at formation generally increase with surface tension
and orifice diameter. Prediction of sizes in swarms from multiple ori-
fices is difficult. In aqueous solutions of low surface tension, bubble
diameters of the order of 1 mm are common. Bubbles produced by
the more complicated techniques of pressure flotation or vacuum
flotation are usually smaller, with diameters of the order of 0.1 mm
Enriching and Stripping Unlike truly simple foam fractiona-
tion without significant changes in bubble diameter, coalescence in
a foam column destroys some bubble surface and so releases
adsorbed material to trickle down through the rising foam. This
downflow constitutes internal reflux, which enriches the rising
foam by countercurrent action. The result is a richer foamate, i.e.,
higher CQ than that obtainable from the single theoretical stage of
the corresponding simple mode. Significant coalescence is often
present in rising foam, but the effect on bubble diameter and
enrichment is frequently overlooked.
External reflux can be furnished by returning some of the externally
broken foam to the top of the column. The concentrating effect of
reflux, even for a substance which saturates the surface, has been
verified [Lemlich and Lavi, Science, 134, 191 (1961)].
Introducing the feed into the foam some distance above the pool
makes for stripping operation. The resulting countercurrent flow in
the foam further purifies the bottoms, i.e., lowers CW.
Enriching, stripping, and combined operations are shown in Fig.
Foam-Column Theory The counterflowing streams within the
foam are viewed as consisting effectively of a descending stream of
interstitial liquid (equal to zero for the simple mode) and an ascend-
ing stream of interstitial liquid plus bubble surface. (By considering
this ascending surface as analogous to a vapor, the overall operation
becomes analogous in a way to distillation with entrainment.)
An effective concentration [Cෆ] in the ascending stream at any level
in the column is defined by Eq. (20-44):
Cෆ = C + (GSΓ/U) (20-44)
where U is the volumetric rate of interstitial liquid upflow, C is the
concentration in this ascending liquid at that level, and Γ is the surface
excess in equilibrium with C. Any effect of micelles should be
For simplicity, U can usually be equated to Q. An effective equi-
librium curve can now be plotted from Eq. (20-44) in terms of Cෆ (or
rather Cෆ°) versus C.
Operating lines can be found in the usual way from material bal-
ances. The slope of each such line is ∆Cෆ/∆C = L/U, where L is the
downflow rate in the particular column section and C is now the con-
centration in the descending stream.
The number of theoretical stages can then be found in one of the
usual ways. Figure 20-42 illustrates a graphical calculation for a stripper.
Alternatively, the number of transfer units (NTU) in the foam based
on, say, the ascending stream can be found from Eq. (20-45):
NTU = ͵
Cෆ° is related to C by the effective equilibrium curve, and Cෆ°W is
similarly related to CW. Cෆ is related to C by the operating line.
To illustrate this integration analytically, Eq. (20-45) becomes Eq.
(20-46) for the case of a stripping column removing a colligend which
is subject to the linear-equilibrium isotherm Γ = KC.
NTU = ln (20-46)
FW + F(GSK − W)CF /CW
GSK(GSK + F − W)
GSK − W
20-32 ALTERNATIVE SEPARATION PROCESSES
As another illustration, Eq. (20-45) becomes Eq. (20-47) for an enrich-
ing column which is concentrating a surfactant with a constant Γ:
NTU = R ln (20-47)
Unless the liquid pool is purposely lengthened vertically in order to
give additional separation via bubble fractionation, it is usually taken
to represent one theoretical stage. A bubbler submergence of 30 cm
or so is usually ample for a solute with a molecular weight that does
not exceed several hundred.
In a colligend stripper, it may be necessary to add some collector to
the pool as well as the feed because the collector is also stripped off.
Limiting Equations If the height of a foam-fractionation col-
umn is increased sufficiently, a concentration pinch will develop
between the counterflowing interstitial streams (Brunner and Lem-
lich, loc. cit.). For an enricher, the separation attained will then
approach the predictions of Eq. (20-48) and, interestingly enough,
CD = CW + (GSΓW /D) (20-48)
D is the volumetric rate at which net foamate (net overhead liquid
product) is withdrawn. D = Q/(R + 1). The concentration in the net
foamate is CD. In the usual case of total foam breakage (no dephleg-
mation), CD = CQ.
If the tall column is a stripper, the separation will approach that of
Eqs. (20-49) and (20-50):
CQ = CF + (GSΓF /Q) (20-49)
CW = CF − (GSΓF /W) (20-50)
For a sufficiently tall combined column, the separation will
approach that of Eqs. (20-51) and (20-50):
CD = CF + (GSΓF /D) (20-51)
The formation of micelles in the foam breaker does not affect the
limiting equations because of the theoretically unlimited opportunity
in a sufficiently tall column for their transfer from the reflux to the
ascending stream [Lemlich, “Principles of Foam Fractionation,” in
Perry (ed.), Progress in Separation and Purification, vol. 1, Inter-
science, New York, 1968, chap. 1].
In practice, the performance of a well-operated foam column sev-
eral feet tall may actually approximate the limiting equations, pro-
vided there is little channeling in the foam and provided that reflux is
either absent or is present at a low ratio.
RGSΓ(F − D)
(R + 1)GSΓ(F − D) − (R + 1)FD(CD − CF)
Column Operation To assure intimate contact between the coun-
terflowing interstitial streams, the volume fraction of liquid in the foam
should be kept below about 10 percent—and the lower the better. Also,
rather uniform bubble sizes are desirable. The foam bubbles will thus
pack together as blunted polyhedra rather than as spheres, and the suc-
tion in the capillaries (Plateau borders) so formed will promote good
liquid distribution and contact. To allow for this desirable deviation
from sphericity, S = 6.3/d in the equations for enriching, stripping, and
combined column operation [Lemlich, Chem. Eng., 75(27), 95 (1968);
76(6), 5 (1969)]. Diameter d still refers to the sphere.
Visible channeling or significant deviations from plug flow of the
foam should be avoided, if necessary by widening the column or low-
ering the gas and/or liquid rates. The superficial gas velocity should
probably not exceed 1 or 2 cm/s. Under proper conditions, HTU val-
ues of several cm have been reported [Hastings, Ph.D. dissertation,
Michigan State University, East Lansing, 1967; and Jashnani and
Lemlich, Ind. Eng. Chem. Process Des. Dev., 12, 312 (1973)]. The
foam column height equals NTU × HTU.
For columns that are wider than several centimeters, reflux and
feed distributors should be used, particularly for wet foam [Haas and
Johnson, Am. Inst. Chem. Eng. J., 11, 319 (1965)]. Liquid content
within the foam can be monitored conductometrically [Chang and
Lemlich, J. Colloid Interface Sci., 73, 224 (1980)]. See Fig. 20-43.
Theoretically, as the limit Ᏸ = K = 0 is very closely approached, Ᏸ = 3K
[Lemlich, J. Colloid Interface Sci., 64, 107 (1978)].
Wet foam can be handled in a bubble-cap column (Wace and Ban-
field, Chem. Process Eng., 47(10), 70 (1966)] or in a sieve plate col-
umn [Aguayo and Lemlich, Ind. Eng. Chem. Process Des. Dev., 13,
153 (1974)]. Alternatively, individual short columns can be connected
in countercurrent array [Banfield, Newson, and Alder, Am. Inst.
Chem. Eng. Symp. Ser., 1, 3 (1965); Leonard and Blacyki, Ind. Eng.
Chem. Process Des. Dev., 17, 358 (1978)].
A high gas rate can be used to achieve maximum throughput in the
simple mode (Wace, Alder, and Banfield, AERE-R5920, U.K. Atomic
Energy Authority, 1968) because channeling is not a factor in that
mode. A horizontal drainage section can be used overhead [Haas and
Johnson, Ind. Eng. Chem. Fundam., 6, 225 (1967)]. The highly mobile
dispersion produced by a very high gas rate is not a true foam but is
ALTERNATIVE SOLID/LIQUID SEPARATIONS 20-33
FIG. 20-42 Graphical determination of theoretical stages for a foam-
fractionation stripping column.
FIG. 20-43 Empirical relationship between Ᏸ, the volumetric fraction of liq-
uid in common polydisperse foam, and K, the electrical conductivity of the foam
divided by the electrical conductivity of the liquid. [Chang and Lemlich, J. Col-
loid Interface Sci., 73, 224 (1980).]
rather a so-called gas emulsion [Bikerman, Ind. Eng. Chem., 57(1),
A very low gas rate in a column several feet tall with internal reflux
can sometimes be used to effect difficult multicomponent separations
in batch operation [Lemlich, “Principles of Foam Fractionation,” in
Perry (ed.), Progress in Separation and Purification, vol. 1, Inter-
science, New York, 1968, chap. 1].
The same end may be achieved by continuous operation at total
external reflux with a small U bend in the reflux line for foamate
holdup [Rubin and Melech, Can. J. Chem. Eng., 50, 748 (1972)].
The slowly rising foam in a tall column can be employed as the sor-
bent for continuous chromatographic separations [Talman and Rubin,
Sep. Sci., 11, 509 (1976)]. Low gas rates are also employed in short
columns to produce the scumlike froth of batch-operated ion flota-
tion, microflotation, and precipitate flotation.
Foam Drainage and Overflow The rate of foam overflow on a
gas-free basis (i.e., the total volumetric foamate rate Q) can be esti-
mated from a detailed theory for foam drainage [Leonard and Lem-
lich, Am. Inst. Chem. Eng. J., 11, 18 (1965)]. From the resulting
relationship for overflow [Fanlo and Lemlich, Am. Inst. Chem. Eng.
Symp. Ser., 9, 75, 85 (1965)], Eq. (20-52) can be employed as a con-
venient approximation to the theory so as to avoid trial and error over
the usual range of interest for foam of low liquid content ascending in
The superficial gas velocity vg is G/A, where A is the horizontal
cross-sectional area of the empty vertical foam column. Also, g is the
acceleration of gravity, ρ is the liquid density, µ is the ordinary liquid
viscosity, and µs is the effective surface viscosity.
To account for inhomogeneity in bubble sizes, d in Eq. (20-52)
should be taken as ͙Σෆnෆidෆi
/ෆΣෆnෆidෆiෆ and evaluated at the top of the verti-
cal column if coalescence is significant in the rising foam. Note that
this average d for overflow differs from that employed earlier for S.
Also, see “Bubble Sizes” regarding the correction for planar statistical
sampling bias and the presence of size segregation at a wall.
For theoretical reasons, Q determined from Eq. (20-52) should be
multiplied by the factor (1 + 3Q/G) to give a final Q. However, for
foam of sufficiently low liquid content this multiplication can be omit-
ted with little error.
The effective surface viscosity is best found by experiment with the
system in question, followed by back calculation through Eq. (20-52).
From the precursors to Eq. (20-52), such experiments have yielded
values of µs on the order of 10−4
(dyn·s)/cm for common surfactants in
water at room temperature, which agrees with independent measure-
ments [Lemlich, Chem. Eng. Sci., 23, 932 (1968); and Shih and Lem-
lich, Am. Inst. Chem. Eng. J., 13, 751 (1967)]. However, the expected
high µs for aqueous solutions of such skin-forming substances as
saponin and albumin was not attained, perhaps because of their non-
newtonian surface behavior [Shih and Lemlich, Ind. Eng. Chem. Fun-
dam., 10, 254 (1971); and Jashnani and Lemlich, J. Colloid Interface
Sci., 46, 13 (1974)].
The drainage theory breaks down for columns with tortuous cross
section, large slugs of gas, or heavy coalescence in the rising foam.
Foam Coalescence Coalescence is of two types. The first is the
growth of the larger foam bubbles at the expense of the smaller bub-
bles due to interbubble gas diffusion, which results from the smaller
bubbles having somewhat higher internal pressures (Adamson and
Gast, Physical Chemistry of Surfaces, 6th ed., Wiley, New York, 1997).
Small bubbles can even disappear entirely. In principle, the rate at
which this type of coalescence proceeds can be estimated [Ranadive
and Lemlich, J. Colloid Interface Sci., 70, 392 (1979)].
The second type of coalescence arises from the rupture of films
between adjacent bubbles [Vrij and Overbeek, J. Am. Chem. Soc., 90,
3074 (1968)]. Its rate appears to follow first-order reaction kinetics
with respect to the number of bubbles [New, Proc. 4th Int. Congr.
Surf. Active Substances, Brussels, 1964, 2, 1167 (1967)] and to de-
crease with film thickness [Steiner, Hunkeler, and Hartland, Trans.
Inst. Chem. Eng., 55, 153 (1977)]. Many factors are involved [Biker-
man, Foams, Springer-Verlag, New York, 1973; and Akers (ed.),
Foams, Academic, New York, 1976].
Both types of coalescence can be important in the foam separations
characterized by low gas flow rate, such as batchwise ion flotation pro-
ducing a scum-bearing froth of comparatively long residence time. On
the other hand, with the relatively higher gas flow rate of foam frac-
tionation, the residence time may be too short for the first type to be
important, and if the foam is sufficiently stable, even the second type
of coalescence may be unimportant.
Unlike the case for Eq. (20-52), when coalescence is significant, it is
better to find S from d evaluated at the feed level for Eqs. (20-49) to
(20-51) and at the pool surface for Eqs. (20-43) and (20-48).
Foam Breaking It is usually desirable to collapse the overflow-
ing foam. This can be accomplished by chemical means (Bikerman,
op. cit.) if external reflux is not employed or by thermal means [Kishi-
moto, Kolloid Z., 192, 66 (1963)] if degradation of the overhead prod-
uct is not a factor.
Foam can also be broken with a rotating perforated basket [Lem-
lich, “Principles of Foam Fractionation,” in Perry (ed.), Progress in
Separation and Purification, vol. 1, Interscience, New York, 1968,
chap. 1]. If the foamate is aqueous (as it usually is), the operation can
be improved by discharging onto Teflon instead of glass [Haas and
Johnson, Am. Inst. Chem. Eng. J., 11, 319 (1965)]. A turbine can be
used to break foam [Ng, Mueller, and Walden, Can. J. Chem. Eng.,
55, 439 (1977)]. Foam which is not overly stable has been broken by
running foamate onto it [Brunner and Stephan, Ind. Eng. Chem.,
57(5), 40 (1965)]. Foam can also be broken by sound or ultrasound, a
rotating disk, and other means [Ohkawa, Sakagama, Sakai, Futai, and
Takahara, J. Ferment. Technol., 56, 428, 532 (1978)].
If desired, dephlegmation (partial collapse of the foam to give reflux)
can be accomplished by simply widening the top of the column, pro-
vided the foam is not too stable. Otherwise, one of the more positive
methods of foam breaking can be employed to achieve dephlegmation.
Bubble Fractionation Figure 20-44 shows continuous bubble
fractionation. This operation can be analyzed in a simplified way in
terms of the adsorbed carry-up, which furthers the concentration gra-
dient, and the dispersion in the liquid, which reduces the gradient
[Lemlich, Am. Inst. Chem. Eng. J., 12, 802 (1966); 13, 1017 (1967)].
To illustrate, consider the limiting case in which the feed stream
and the two liquid takeoff streams of Fig. 20-44 are each zero, thus
resulting in batch operation. At steady state the rate of adsorbed carry-
up will equal the rate of downward dispersion, or afΓ = DෆAdC/dh.
Here a is the surface area of a bubble, f is the frequency of bubble for-
mation. Dෆ is the dispersion (effective diffusion) coefficient based on
the column cross-sectional area A, and C is the concentration at height
h within the column.
20-34 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-44 Continuous bubble fractionation.
There are several possible alternative relationships for Γ (Lemlich,
op. cit.). For simplicity, consider Γ = K′C, where K′ is not necessarily
the same as the equilibrium constant K. Substituting and integrating
from the boundary condition of C = CB at h = 0 yield
C/CB = exp (Jh) (20-53)
CB is the concentration at the bottom of the column, and parameter
J = K′af/DෆA. Combining Eq. (20-53) with a material balance against
the solute in the initial charge of liquid gives
Ci is the concentration in the initial charge, and H is the total height of
The foregoing approach has been extended to steady continuous
flow as illustrated in Fig. 20-44 [Cannon and Lemlich, Chem. Eng.
Prog. Symp. Ser., 68(124), 180 (1972); Bruin, Hudson, and Morgan,
Ind. Eng. Chem. Fundam., 11, 175 (1972); and Wang, Granstrom, and
Kown, Environ. Lett., 3, 251 (1972), 4, 233 (1973), 5, 71 (1973)]. The
extension includes a rough method for estimating the optimum feed
location as well as a very detailed analysis of column performance
which takes into account the various local phenomena around each
rising bubble (Cannon and Lemlich, op. cit.).
Uraizee and Narsimhan [Sep. Sci. Technol., 30(6), 847 (1995)] have
provided a model for the continuous separation of proteins from
dilute solutions. Although their work is focused on protein separation,
the model should find general application to other separations.
In agreement with experiment [Shah and Lemlich, Ind. Eng.
Chem. Fundam., 9, 350 (1970); and Garmendia, Perez, and Katz, J.
Chem. Educ., 50, 864 (1973)], theory shows that the degree of sepa-
ration that is obtained increases as the liquid column is made taller.
But unfortunately it decreases as the column is made wider. In simple
terms, the latter effect can be attributed to the increase in the disper-
sion coefficient as the column is widened.
In this last connection it is important that the column be aligned
precisely vertically (Valdes-Krieg, King, and Sephton, Am. Inst.
Chem. Eng. J., 21, 400 (1975)]. Otherwise, the bubbles with their
dragged liquid will tend to rise up one side of the column, thus caus-
ing liquid to flow down the other side, and in this way largely destroy
the concentration gradient. A vertical foam-fractionation column
should also be carefully aligned to be plumb.
The escaping bubbles from the top of a bubble-fractionation col-
umn can carry off an appreciable quantity of adsorbed material in
an aerosol of very fine film drops [various papers, J. Geophys. Res.,
Oceans Atmos., 77(27), (1972)]. If the residual solute is thus appre-
ciably depleted, Ci in Eq. (20-54) should be replaced with the average
This carry-off of film drops, which may also occur with breaking
foam, in certain cases can partially convert water pollution into air pol-
lution. If such is the case, it may be desirable to recirculate the gas.
Such recirculation is also indicated if hydrocarbon vapors or other
volatiles are incorporated in the gas stream to improve adsorptive
selectivity [Maas, Sep. Sci., 4, 457 (1969)].
A small amount of collector (surfactant) or other appropriate addi-
tive in the liquid may greatly increase adsorption (Shah and Lemlich,
op. cit.). Column performance can also be improved by skimming the
surface of the liquid pool or, when possible, by removing adsorbed
solute in even a tenuous foam overflow. Alternatively, an immiscible
liquid can be floated on top. Then the concentration gradient in the
tall pool of main liquid, plus the trapping action of the immiscible
layer above it, will yield a combination of bubble fractionation and sol-
Systems Separated Some of the various separations reported in
the literature are listed in Rubin and Gaden, “Foam Separation,” in
Schoen (ed.), New Chemical Engineering Separation Techniques, Inter-
science, New York, 1962, chap. 5; Lemlich, Ind. Eng. Chem., 60(10), 16
(1968); Pushkarev, Egorov, and Khrustalev, Clarification and Deactiva-
tion of Waste Waters by Frothing Flotation, in Russian, Atomizdat,
Moscow, 1969; Kuskin and Golman, Flotation of Ions and Molecules, in
Russian, Nedra, Moscow, 1971; Lemlich (ed.), Adsorptive Bubble
JH exp (Jh)
exp (JH) − 1
Separation Techniques, Academic, New York, 1972; Lemlich, “Adsub-
ble Methods,” in Li (ed.), Recent Developments in Separation Science,
vol. 1, CRC Press, Cleveland, 1972, chap. 5; Grieves, Chem. Eng. J., 9,
93 (1975); Valdes-Krieg King, and Sephton, Sep. Purif. Methods, 6, 221
(1977); Clarke and Wilson, Foam Flotation, Marcel Dekker, New York,
1983; and Wilson and Clarke, “Bubble and Foam Separations in Waste
Treatment,” in Rousseau (ed.), Handbook of Separation Processes,
Wiley, New York, 1987.
Of the numerous separations reported, only a few can be listed
here. Except for minerals beneficiation [ore flotation], the most
important industrial applications are usually in the area of pollution
A pilot-sized foaming unit reduced the alkyl benzene sulfonate con-
centration of 500,000 gal of sewage per day to nearly 1 mg/L, using a
G/F of 5 and producing a Q/F of no more than 0.03 [Brunner and
Stephan, Ind. Eng. Chem., 57(5), 40 (1965); and Stephan, Civ. Eng.,
35(9), 46 (1965)]. A full-scale unit handling over 45,420 m3
million gal/day) performed nearly as well. The foam also carried off
some other pollutants. However, with the widespread advent of
biodegradable detergents, large-scale foam fractionation of municipal
sewage has been discontinued.
Other plant-scale applications to pollution control include the flota-
tion of suspended sewage particles by depressurizing so as to release
dissolved air [Jenkins, Scherfig, and Eckhoff, “Applications of Adsorp-
tive Bubble Separation Techniques to Wastewater Treatment,” in
Lemlich (ed.), Adsorptive Bubble Separation Techniques, Academic,
New York, 1972, chap. 14; and Richter, Internat. Chem. Eng., 16, 614
(1976)]. Dissolved-air flotation is also employed in treating waste-
water from pulp and paper mills [Coertze, Prog. Water Technol., 10,
449 (1978); and Severeid, TAPPI 62(2), 61, 1979]. In addition, there is
the flotation, with electrolytically released bubbles [Chambers and
Cottrell, Chem. Eng., 83(16), 95 (1976)], of oily iron dust [Ellwood,
Chem. Eng., 75(16), 82 (1968)] and of a variety of wastes from sur-
face-treatment processes at the maintenance and overhaul base of an
airline [Roth and Ferguson, Desalination, 23, 49 (1977)].
Fats and, through the use of lignosulfonic acid, proteins can be
flotated from the wastewaters of slaughterhouses and other food-
processing installations [Hopwood, Inst. Chem. Eng. Symp. Ser., 41,
M1 (1975)]. After further treatment, the floated sludge has been fed
A report of the recovery of protein from potato-juice wastewater by
foaming [Weijenberg, Mulder, Drinkenberg, and Stemerding, Ind.
Eng. Chem. Process Des. Dev., 17, 209 (1978)] is reminiscent of the
classical recovery of protein from potato and sugar-beet juices [Ost-
wald and Siehr, Kolloid Z., 79, 11 (1937)]. The isoelectric pH is often
a good choice for the foam fractionation of protein (Rubin and Gaden,
loc. cit.). Adding a salt to lower solubility may also help. Additional
applications of foam fractionation to the separation of protein have
been reported by Uraizee and Narsimhan [Enzyme Microb. Technol.
12, 232 (1990)].
With the addition of appropriate additives as needed, the flotation
of refinery wastewaters reduced their oil content to less than 10 mg/L
in pilot-plant operation [Steiner, Bennett, Mohler, and Clere, Chem.
Eng. Prog., 74(12), 39 (1978)] and full-scale operation (Simonsen,
Hydrocarb. Process. Pet. Refiner, 41(5), 145, 1962]. Experiments with
a cationic collector to remove oils reportedly confirmed theory
[Angelidon, Keskavarz, Richardson, and Jameson, Ind. Eng. Chem.
Process Des. Dev., 16, 436 (1977)].
Pilot-plant [Hyde, Miller, Packham, and Richards, J. Am. Water
Works Assoc., 69, 369 (1977)] and full-scale [Ward, Water Serv., 81,
499 (1977)] flotation in the preparation of potable water is described.
Overflow at the rate of 2700 m3
(713,000 gal) per day from a zinc-
concentrate thickener is treated by ion flotation, precipitate flotation,
and untrafine-particle flotation [Nagahama, Can. Min. Metall. Bull.,
67, 79 (1974)]. In precipitate flotation only the surface of the particles
need be coated with collector. Therefore, in principle less collector is
required than for the equivalent removal of ions by foam fractionation
or ion flotation.
By using an anionic collector and external reflux in a combined
(enriching and stripping) column of 3.8-cm (1.5-in) diameter with a
feed rate of 1.63 m/h [40 gal/(h⋅ft2
)] based on column cross section,
ALTERNATIVE SOLID/LIQUID SEPARATIONS 20-35
D/F was reduced to 0.00027 with Cw /CF for Sr2+
below 0.001 [Shon-
feld and Kibbey, Nucl. Appl., 3, 353 (1967)]. Reports of the adsubble
separation of 29 heavy metals, radioactive and otherwise, have been
tabulated [Lemlich, “The Adsorptive Bubble Separation Techniques,”
in Sabadell (ed.), Proc. Conf. Traces Heavy Met. Water, 211–223,
Princeton University, 1973, EPA 902/9-74-001, U.S. EPA, Reg. II,
1974). Some separation of 15
N from 14
N by foam fractionation has
been reported [Hitchcock, Ph.D. dissertation, University of Missouri,
The numerous separations reported in the literature include sur-
factants, inorganic ions, enzymes, other proteins, other organics, bio-
logical cells, and various other particles and substances. The scale of
the systems ranges from the simple Crits test for the presence of sur-
factants in water, which has been shown to operate by virtue of tran-
sient foam fractionation [Lemlich, J. Colloid Interface Sci., 37, 497
(1971)], to the natural adsubble processes that occur on a grand scale
in the ocean [Wallace and Duce, Deep Sea Res., 25, 827 (1978)]. For
further information see the reviews cited earlier.
20-36 ALTERNATIVE SEPARATION PROCESSES
MEMBRANE SEPARATION PROCESSES
GENERAL REFERENCES: Cheryan, Ultrafiltration and Microfiltration Hand-
book, Technomic Publishing Company, Pa., 1988. Ho and Sirkar (eds.), Mem-
brane Handbook, Van Nostrand Reinhold, New York, 1992. Baker, Membrane
Technology and Applications, 2d ed., Wiley, 2000. Eykamp, Sec. 22: Mem-
brane Separation Processes, in Perry’s Chemical Engineers’ Handbook, 7th
ed., Perry and Green (eds.), McGraw-Hill, New York, 1997. Millipore Corpo-
ration, Protein Concentration and Diafiltration by Tangential Flow Filtration,
Lit. No. TB032 Rev. B, 1999. Zeman and Zydney, Microfiltration and Ultrafil-
tration: Principles and Applications, Marcel Dekker, New York, 1996. A
review of membrane retention mechanisms can be found in Deen, AIChE J.,
33, 1409–1425 (1987).
New developments in membranes are found in journals and trade
magazines (e.g., J. Membrane Sci., BioPharm International), vendor
communications (e.g., websites), patent filings, and conference pre-
sentations (e.g., annual ACS or NAMS meetings). Areas of active
research include new membrane polymers and surface modification
with accompanying diagnostic methods (to reduce fouling, increase
flux and retention, improve consistency), new module designs (to
improve flux, cleanability, ease of use, scalability, reliability), new
processing skids (better components, recovery, less holdup, better
mixing, disposability, software for automated processing and archiv-
ing), new processing methods (diafiltration strategies, turbulence
enhancements), and new applications (e.g., protein-protein separa-
Topics Omitted from This Section In order to concentrate on the
membrane processes of widest industrial interest, several are left out.
Dialysis and Hemodialysis Historically, dialysis has found some
industrial use. Today, much of that is supplanted by ultrafiltration.
Donan dialysis is treated briefly under electrodialysis. Hemodialysis is
a huge application for membranes, and it dominates the membrane
field in area produced and in monetary value. This medical applica-
tion is omitted here.
An excellent description of the engineering side of both topics is
provided by Kessler and Klein [in Ho and Sirkar (eds.), op. cit., pp.
163–216]. A comprehensive treatment of diffusion appears in: Von
Halle and Shachter, “Diffusional Separation Methods,” in Encyclope-
dia of Chemical Technology, pp. 149–203, Wiley, 1993.
Facilitated Transport Transport by a reactive phase through a
membrane is promising but problematic. Way and Noble [in Ho and
Sirkar (eds.), op. cit., pp. 833–866] have a description and a complete
Liquid Membranes Several types of liquid membranes exist:
molten salt, emulsion, immobilized/supported, and hollow-fiber-
contained liquid membranes. Araki and Tsukube (Liquid Membranes:
Chemical Applications, CRC Press, 1990) and Sec. IX and Chap. 42 in
Ho and Sirkar (eds.) (op. cit., pp. 724, 764–808) contain detailed
information and extensive bibliographies.
Catalytic Membranes Falconer, Noble, and Sperry (Chap. 14—
“Catalytic Membrane Reactors” in Noble and Stern, op. cit., p. 669–
712) give a detailed review and an extensive bibliography. Additional
information can be found in a work by Tsotsis et al. [“Catalytic Mem-
brane Reactors,” pp. 471–551, in Becker and Pereira (eds.), Com-
puter-Aided Design of Catalysts, Dekker, 1993].
GENERAL BACKGROUND AND DEFINITIONS
Applications Membranes create a boundary between different
bulk gas or liquid mixtures. Different solutes and solvents flow
through membranes at different rates. This enables the use of mem-
branes in separation processes. Membrane processes can be operated
at moderate temperatures for sensitive components (e.g., food, phar-
maceuticals). Membrane processes also tend to have low relative cap-
ital and energy costs. Their modular format permits reliable scale-up
and operation. This unit operation has seen widespread commercial
adoption since the 1960s for component enrichment, depletion, or
equilibration. Estimates of annual membrane module sales in 2005
are shown in Table 20-16. Applications of membranes for diagnostic
and bench-scale use are not included. Natural biological systems
widely employ membranes to isolate cells, organs, and nuclei.
Common Definitions Membrane processes have been evolving
since the 1960s with each application tending to generate its own ter-
minology. Recommended nomenclature is provided along with alter-
natives in current use.
Fluid Stream Designations For the generalized membrane
module shown in Fig. 20-45, a feed stream enters a membrane
module while both a permeate and a retentate stream exit the mod-
ule. The permeate (or filtrate) stream flows through the membrane
and has been depleted of retained components. The term filtrate is
commonly used for NFF operation while permeate is used for TFF
operation. The retentate (or concentrate) stream flows through the
module, not the membrane, and has been enriched in retained
TABLE 20-16 Membrane Market in 2005
Segment $M/yr Size Applications Characteristics
Dialysis ~2,000 Medical Mature growing 5%
Reverse osmosis ~500 Water treatment Growing 10%
Microfiltration ~500 Water, food, pharm.
Ultrafiltration ~400 Water, food, pharm. Growing 10%
Gas separation ~500 Nitrogen
Electrodialysis ~100 Water
Pervaporation ~5 Solvent/water Nascent
Facilitated transport 0 None In development
Flow Types: Normal Flow and Tangential Flow Filtration If
the retentate flow in Fig. 20-45 is zero and all the feed stream flows as
a velocity vector normal to the membrane surface, this type of filtra-
tion is referred to as normal flow filtration (NFF), also called dead-
end flow. If there is a stream that flows as a velocity vector tangent to
the membrane surface, creating a velocity gradient at the membrane
surface, and that exits the module as a retentate stream, this is
referred to as tangential flow filtration (TFF), also called crossflow fil-
tration. A retentate stream that flows through a module without creat-
ing a surface velocity gradient is merely a bypass and not TFF.
TFF mitigates the accumulation of retained components on the
membrane surface, reducing the plugging of the membrane and per-
mitting a more steady-state operation. TFF is used in processing feeds
with a high concentration of retained components.
Flow: Flux, Permeability, Conversion The productivity of a
membrane module is described by its flux J = volumetric permeate
flow rate/membrane area with units of volume per area per time. Rel-
atively high flux rates imply that relatively small membrane areas are
required. The permeate volume is usually greater than the feed vol-
ume for a given process. Flux is also the magnitude of the normal flow
velocity with units of distance per time.
The sensitivity of productivity or flux to transmembrane pressure
(TMP) is referred to as the permeability L = flux/transmembrane pres-
sure. TMP refers to a module average. Pure-component permeability
(e.g., water permeability) refers to membrane properties while the
more industrially relevant process permeability includes fouling and
The efficiency of a membrane module is characterized by the recov-
ery or conversion ratio CR = permeate flow rate/feed flow rate. Low
conversion means that fluid has to be repeatedly cycled past the TFF
module to generate permeate. High-efficiency NFF has CR = 1.
Flux Decline: Plugging, Fouling, Polarization Membranes
operated in NFF mode tend to show a steady flux decline while those
operated in TFF mode tend to show a more stable flux after a short
initial decline. Irreversible flux decline can occur by membrane com-
pression or retentate channel spacers blinding off the membrane.
Flux decline by fouling mechanisms (molecular adsorption, precipita-
tion on the membrane surface, entrapment within the membrane
structure) are amenable to chemical cleaning between batches. Flux
decline amenable to mechanical disturbance (such as TFF operation)
includes the formation of a secondary structure on the membrane sur-
face such as a static cake or a fluid region of high component concen-
tration called a polarization layer.
Understanding polarization and controlling its effects are key to
implementing a good TFF process. Solutes entrained by the permeate
flow are retained by the membrane. They accumulate on the mem-
brane surface and form a region of high concentration called the
polarization boundary layer. A steady state is reached between back
transport away from the membrane surface, tangential convective
transport along the membrane surface, and normal convective flow
toward the membrane. The back transport leading to steady-state
operation gives TFF a high capacity. Plugging is commonly used to
describe flux decline through fouling and caking mechanisms.
Single-Component Separation: Passage, Retention, LRV The
passage of a component through a membrane (also called the sieving
coefficient or transmission) is calculated as the ratio S = cP/cF, where
cP and cF refer to the permeate concentration, and the feed concen-
tration, respectively. These concentrations may change with position
within a module. In industrial practice, it is common to measure these
concentrations in the permeate stream exiting a module and the feed
stream entering a module to give an observed passage or Sobs.
Membrane vendors and researchers may use the feed concentra-
tion at the membrane surface or cw to calculate an intrinsic passage
Sint, = cP/cw. The intrinsic passage characterizes the membrane while
the observed passage characterizes the module.
Complementary to the passage, one can also consider the retention of
a component as R = 1 − S (also called rejection). Retention can also be
either an observed or an intrinsic measurement. Retention is useful in
considering retained products during concentration mode operation.
Other component separation characterizations include the log reduc-
tion value LRV = − logS which is used to characterize high-efficiency
separations with permeate products (sterilization). The beta ratio β =
1/S is sometimes used in NFF for clarification applications.
Multiple-Component Separation: Separation Factor Consis-
tent with the characterization of different separation methods, one
can define a separation factor αij (also called selectivity) for compo-
nents i and j that compares their relative concentrations in the perme-
ate stream to those in the feed stream:
αij = ᎏ
ᎏ = ᎏ
The convention is to designate the component with higher passage
as component i so that ␣ij > 1.
Membrane Types Key membrane properties include their size
rating, selectivity, permeability, mechanical robustness (to allow mod-
ule fabrication and withstand operating conditions), chemical robust-
ness (to fabrication materials, process fluids, cleaners, and sanitizers),
low extractibles, low fouling characteristics, high capacity, low cost,
Size Ratings The relative sizes of common components and the
associated membrane classes capable of retaining them are shown in
Vendors characterize their filters with ratings indicating the approx-
imate size (or corresponding molecular weight) of components
retained by the membrane. This rating should be used as a rough
guide only and followed up with retention testing. Among the factors
affecting retention are the application-specific retention require-
ments, variable component size and shapes depending on solution
environment, membrane fouling and compaction, degree of module
polarization, and interaction between feed components.
Composition and Structure Commercial membranes consist
primarily of polymers and some ceramics. Other membrane types
include sintered metal, glass, and liquid film. Polymeric membranes
are formed by precipitating a 5 to 25 wt % casting solution (lacquer)
into a film by solvent evaporation (air casting), liquid extraction
(immersion casting), or cooling (melt casting or thermally induced
phase separation). Membranes can be cast as flat sheets on a variety of
supports or as fibers through a die. Ceramic membranes are formed
by depositing successive layers of smaller and smaller inorganic parti-
cles on a monolith substrate, to create smaller and smaller interstices
or “pores” between particles. These layers are then sintered below the
melting point to create a rigid structure.
Membranes may be surface-modified to reduce fouling or improve
chemical resistance. This can involve adding surface-modifying agents
directly to the lacquer or modifying the cast membrane through
chemical or physical treatment. Membranes can also be formed by
selective etching or track-etching (radiation treatment followed by
etching). Stretching is used to change pore morphology.
Liquid film membranes consist of immiscible solutions held in
membrane supports by capillary forces. The chemical composition of
these solutions is designed to enhance transport rates of selected com-
ponents through them by solubility or coupled chemical reaction.
Membrane Morphology—Pores, Symmetric, Composite Only
nucleopore and anodyne membranes have relatively uniform pores.
Reverse osmosis, gas permeation, and pervaporation membranes
have nonuniform angstrom-sized pores corresponding to spaces in
between the rigid or dynamic membrane molecules. Solute-membrane
molecular interactions are very high. Ultrafiltration membranes have
nonuniform nanometer sized pores with some solute-membrane inter-
actions. For other microfiltration membranes with nonuniform pores
on the submicrometer to micrometer range, solute-membrane interac-
tions are small.
MEMBRANE SEPARATION PROCESSES 20-37
FIG. 20-45 Fluid stream schematic.
Permeate or filtrate
Fluid density and component brownian diffusivity D are also assumed
constant. A steady-state component mass balance can be written for
component concentration c:
u + v = D u and v are normal and tangential velocities
Further neglecting the first term allows integration from y = 0 at
the wall (membrane surface) into the boundary layer. At the wall, the
net flux is represented by convection into the permeate
vc − D = vcp (20-57)
This can be further integrated from the wall to the boundary layer
thickness y = δ, where the component is at the bulk concentration cb.
Substituting J = − v and k = D/δ, the mass-transfer coefficient yields
the stagnant film model [Brian, Desalination by Reverse Osmosis,
Merten (ed.), M.I.T. Press, Cambridge, Mass., 1966, pp. 161–292]:
= − ᎏ
Although allowing for axial variations in v with a constant wall con-
centration cw yields Eq. (20-59), as a more rigorous expression applic-
able to higher concentrations [Trettin et al., Chem. Eng. Comm., 4,
507 (1980)] the form of Eq. (20-58) is convenient in analyzing a vari-
ety of complex behavior.
J = 2.046k
Substituting the intrinsic passage into Eq. (20-58) and rearranging
yields an expression for the polarization modulus cw/cb:
ᎏ = ᎏ
1 + Sin
Figure 20-48 shows Wijmans’s plot [Wijmans et al., J. Membr. Sci.,
109, 135 (1996)] along with regions where different membrane
processes operate (Baker, Membrane Technology and Applications, 2d
ed., Wiley, 2004, p. 177). For RO and UF applications, Sint < 1, and
cw > cb. This may cause precipitation, fouling, or product denatura-
tion. For gas separation and pervaporation, Sint >1 and cw < cb. MF is
not shown since other transport mechanisms besides Brownian diffu-
sion are at work.
cb − cp
cw − cp
Substituting the observed passage into Eq. (20-60) and rearranging
yields Eq. (20-61). A plot of the LHS versus J data yields the mass-
transfer coefficient from the slope, similar to the Wilson plot for heat
A generalized model of transport allowing for component interac-
tions is provided by nonequilibrium thermodynamics where the flux
of component i through the membrane Ji [gmol/(cm2
·s)] is written
as a first-order perturbation of the chemical potential dµi/dx
and Li [1/(cal⋅cm⋅s)] is a component permeability. The chemical poten-
tial gradient provides the driving force for flow and may include
contributions from pressure, concentration, electrical field, and tem-
perature. A complementary description of transport is obtained by
modeling each transport mechanism in some detail. See each applica-
tion section for appropriate models.
Application of Eq. (20-62) to the polarization for two components.
A and B requires solution of the component balances (Saksena, Ph.D.
dissertation, University of Delaware, 1995):
u + v =
DA + DAB
ᎏ + v =
DB + DAB
Larger components with smaller brownian diffusivity polarize more
readily. They can exclude smaller components, reducing their concen-
tration at the membrane surface and increasing their retention by the
Solvent Transport through the Membrane The fluid flux J
through a membrane is commonly modeled as an assembly of np
parallel tubes per unit area with a uniform radius r and length l dis-
playing Hagen-Poiseuille flow:
J = ∆P (20-64)
where η is the fluid viscosity and ∆P the pressure drop. The depen-
dence of flow on r4
means that larger pores take most of the flow.
MEMBRANE SEPARATION PROCESSES 20-39
1.0E-03 1.0E-02 1.0E-01 1.0E+00 1.0E+01
Peclet Number J/k
FIG. 20-48 TFF polarization.
MODULES AND MEMBRANE SYSTEMS
Membranes are packaged into modules for convenient operation. Key
module properties include mechanical robustness (providing good
sealing between fluid streams and membrane support), good flow dis-
tribution (low pressure drops, high TFF membrane shear, low dead
volumes, and low sensitivity to feed channel plugging), cleanability,
chemical robustness, low extractibles, low cost, ease of assembly, scal-
ability, high product recovery, and high consistency. Cleanability con-
siderations evolved from the dairy industry (3A Sanitary Standards for
Crossflow Membrane Modules, USPHS No. 45-00, 1990). Modules
may be integrity-tested for the absence of punctures, improper seal-
ing, or other defects.
NFF modules (Figs. 20-49 and 20-50) include cartridges or cap-
sules, stacked disks, and flat sheet. Pleated cartridges, the most com-
mon format, are available in a variety of sizes and are installed into
multiround stainless-steel housings. As shown in Fig. 20-49, multiple
layers of materials such as support layers or membranes with different
pore ratings may be pleated together. Capsules are made by sealing
individual cartridges within their own plastic housing. Pleated car-
tridge manufacturing is shown in Fig. 20-51.
TFF module types include plate-and-frame (or cassettes), hollow
fibers, tubes, monoliths, spirals, and vortex flow. Figures 20-52 and
20-53 show several common module types and the flow paths within
each. Hollow fiber or tubular modules are made by potting the cast
membrane fibers or tubes into end caps and enclosing the assembly in
a shell. Similar to fibers or tubes, monoliths have their retentive layer
coated on the inside of tubular flow channels or lumens with a high-
permeability porous structure on the shell side.
Spiral-wound modules are made by assembling a membrane packet
or leaf consisting of a permeate channel spacer sheet with two flat
sheets of membranes on either side (facing outward from the perme-
ate sheet). The sides of this packet are glued together, and the top is
glued to a central core collection tube containing small holes that
allow passage of permeate from the spacer channel into the collection
tube. A feed-side spacer is placed on the packet, and the packet is
wound around the core to create the module. The wound packet
assembly is sometimes inserted into a shell. Spiral modules frequently
use a multileaf construction in which several packets are glued to the
central core before winding. This design reduces the pressure drop
associated with a long permeate path length.
Cassettes use feed and permeate spacers with one set of precut
holes for the feed/retentate and another, smaller set of holes for
the permeate. The feed spacer has a raised-edge seal around the
permeate hole to prevent flow, and the permeate spacer has a simi-
lar seal around the feed/retentate hole. Similar to spiral assembly,
membrane packets are glued together around their edges and
assembled into a stack, and the stack is glued together around the
For fibers or tubular modules, the feed is generally introduced into
the inside of the tubes, or lumen, while permeate is withdrawn from
the shell side. This flow orientation enhances the shear at the mem-
brane surface for TFF operation. These modules may also be run at
high conversion or NFF mode with the feed introduced on the out-
side of the tubes or shell side. In this case, the shell side offers greater
Table 20-17 summarizes TFF module characteristics. The pre-
ferred TFF module type depends on the characteristics of each
In designing TFF modules, the polarization equation [Eq. (20-58)]
indicates that for given fixed concentrations, increases in the mass-
transfer coefficient k will increase the flux. This can be accomplished
by (1) increasing the shear rate at the wall through higher tangential
flow velocities and smaller channel heights, or mechanically moving
the membrane by spinning or vibration; (2) altering the geometry of
the feed channel to increase turbulent mixing in the normal direction
by using feed channel screens or curving channels to introduce Dean
or Taylor vortices; (3) introducing pulsatile flow in the feed channel or
periodic bursts of gas bubbles; and (4) introducing body forces (cen-
trifugal, gravitational, or electromagnetic) to augment transport away
from the membrane and introducing large particles in the feed (e.g.,
>0.5 µm PVC latex beads) that disrupt boundary layers and cause
Another device that finds frequent use is the stirred cell shown in
Fig. 20-54. This device uses a membrane coupon at the bottom of the
reservoir with a magnetic stir bar. Stirred cells use low fluid volumes
and can be used in screening and R&D studies to evaluate membrane
types and membrane properties. The velocity profiles have been well
defined (Schlichting, Boundary Layer Theory, 6th ed., McGraw-Hill,
New York, 1968, pp. 93–99).
20-40 ALTERNATIVE SEPARATION PROCESSES
NONUNIFORM FLOW DISTRIBUTION
AMOUNT OF FLOW DEPICTED
BY SIZE OF ARROW
Conventional Pleated Element Construction
FIG. 20-49 NFF pleated cartridges and capsules. (Courtesy Millipore Corporation, Pall Corporation.)
FIG. 20-50 NFF stacked disk modules. (Courtesy Millipore Corporation.)
Mass-transfer coefficients for a single newtonian component in var-
ious module types and flow regimes can be correlated by Eq. (20-65)
with values for the constants in Table 20-18:
Sh = aReb
where Sh = kD/dh
Re = ᎏ
ᎏ (tube, slit or spacer with hydraulic diameter dh)
Re = ᎏ
ᎏ (stirred cell diameter d)
Re = ᎏ
ᎏ (rotating-cylinder inner radius R and gap diameter d)
Sc = ᎏ
MEMBRANE SEPARATION PROCESSES 20-41
FIG. 20-52 Commercial TFF modules: (a) hollow fibers; (b) spirals; (c) cassettes. (Courtesy Millipore Corporation.)
FIG. 20-51 Pleated cartridge manufacturing.
(a) (b) (c)
Membrane Manufacturing Lamination Pleating
Serial No.Ultrasonic SeamingEnd Cap Welding
100% Integrity Testing Stacking 100% Physical Testing
Lot Release TestingPackaging
B. diminuta Spike
Large-Area Configurations NFF membrane systems consist
primarily of large, multiround housings that contain pleated cartridges
inserted into a common base plate. These cartridges share a common
feed and collect a common filtrate.
TFF membrane systems generally use a common feed distributed
among parallel modules with a collection of common retentate and
common permeate streams. In some applications, it is also useful to
plumb TFF modules with the retentate in series where the retentate
flow from one module provides the feed flow to the next module. This
type of configuration is equivalent to increasing the length of the
retentate channel. Permeate flows may or may not be plumbed
Process Configurations Basic membrane process configura-
tions shown in Fig. 20-55 include single-pass, batch, fed-batch, and
Single-Pass Operation A single-pass configuration is used for
NFF system operation (no retentate). Component concentrations
change along the length of the retentate channel. For dilute streams
such as water, assemblies of TFF modules that run in NFF mode
(retentate closed) are also used. For nondilute streams, a single pass
through a module may be insufficient to generate either the desired
permeate flow or a concentrated retentate.
Steady-state component and solvent mass balances can be written for
single-pass operation by considering an incremental area element dA in
the axial or flow direction for a feed channel, module, or membrane
assembly of constant width. For volumetric crossflow Q in the feed
channel and observed solute passage Si,
= ci + Q = − JciSi (20-66)
ᎏ = − J (20-67)
Combining these equations and integrating yield ci = cio X1−Si
volume reduction factor X = Q/Qo and the observed component pas-
sage Si. This allows one to determine either final concentrations from
crossflow rates or the reverse. For a fully retained product (Si = 0), a
10-fold volume reduction (X = 10) produces a 10-fold more concen-
trated product. However, if the product is only partially retained, the
volume reduction does not proportionately increase the final concen-
tration due to losses through the membrane.
The area required for processing A = (Qo − Q)/J
, where Qo − Q is
the permeate volumetric flow, can be estimated by using the approx-
≈ 0.33Jinitial + 0.67Jfinal (Cheryan, Ultrafiltration Handbook,
Technomic, Lancaster, Pa., 1986) and a suitable flux model. An
appropriate model relating flux to crossflow, concentration, and pres-
sure is then applied. Pressure profiles along the retentate channel are
empirically correlated with flow for spacer-filled channels to obtain
∆P = ∆Po(Q/A)n
20-42 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-53 TFF module flow paths: (a) hollow fibers; (b) spirals; (c) cassettes. (Courtesy Millipore Corporation.)
TABLE 20-17 Commercial TFF Modules
Spiral Fiber Cassette
Screens/spacers Yes No Yes/no
Typical no. in series 1–2 1–2 1–2
Packing density, m2
800 1000–6000 500
Feed flow, L/m2
/h (LMH) 700–5000 500–18,000 400
Feed pressure drop, psi/module 5–15 1–5 10–50
Channel height, mm 0.3–1 0.2–3 0.3–1
Plugging sensitivity High Moderate High
Working volume, L/m2
1 0.5 0.4
Holdup volume, L/m2
0.03 0.03 0.02
Module cost, $/m2
40–200 200–900 500–1000
Ruggedness Moderate Low-moderate High
Module areas, m2
0.1–35 0.001–5 0.05–2.5
Membrane types RO-UF RO-UF-MF UF-MF
Relative mass-transfer efficiency* 6 4 10
Ease of use Moderate High Moderate
Scalability† Fair Moderate Good
*Qualitative, based on relative fluxes.
†Cassettes keep retentate path length constant and require lower feed flow rates.
(a) (b) (c)
Batch and Fed-Batch Operation Batch operation involves
recycling the retentate to the feed tank to create a multipass flow of
the feed through the module. The compositions in batch systems
change over time. Recycling is necessary for nondilute streams to gen-
erate either the desired permeate flow or a concentrated retentate.
Multipass operation increases residence times and pump passes
that may degrade retentate components. For systems requiring high
pressures to generate permeate flow, it is useful to run the entire sys-
tem at high pressure to save energy.
An unsteady-state component mass balance, Eq. (20-68), can be
written for batch operation by assuming a uniform average retentate
concentration ci within the system. Assuming a constant solvent con-
centration and a 100 percent passage, the solvent balance becomes
= − JAciSi (20-68)
= − JA (20-69)
Dividing Eq. (20-68) by Eq. (20-69) and integrating yield the rela-
tionship in Table 20-19 between retentate concentration, the volume
reduction factor X = (initial retentate volume)/(final retentate volume),
and the observed component passage Si. For a fully retained product
(Si =0), a 10-fold volume reduction (X = 10) produces a 10-fold more
concentrated product. However, if the product is only partially
retained, the volume reduction does not proportionately increase the
final concentration due to losses through the membrane. The compo-
nent mass in the retentate is obtained as the product of the retentate
concentration and retentate volume, as shown in Table 20-19.
TFF systems have a maximum volume reduction capability, typi-
cally about 40X. This arises because tanks must be large enough to
hold the batch volume while allowing operation at a minimum work-
ing volume. The minimum working volume may be limited by air
entrainment into the feed pump, mixing in the feed tank, or level
measurement capabilities. Fed-batch operation extends the volume
reduction capability to 100X and provides some flexibility in process-
ing a variety of batch volumes in a single skid. For a fed-batch opera-
tion, the retentate is returned to a smaller tank, not the large feed
tank. Feed is added to the small retentate tank as permeate is with-
drawn, so the system volume remains constant. The smaller retentate
tank can allow a smaller working volume without entrainment. While
one could make the retentate tank a section of pipe that returns the
retentate directly into the pump feed (bypass line), such a configura-
tion can be problematic to flush, vent, and drain, leading to cleaning
and product recovery issues.
Equation (20-70) is the unsteady-state component mass balance for
fed-batch concentration at constant retentate volume. Integration
yields the equations for concentration and yield in Table 20-19.
V = JAcF − JAciSi (20-70)
Retentate concentration over the course of the process is shown in
Fig. 20-56 for a fully retentive product (Si = 0). The tank ratio =
V0 /retentate tank volume, C0 is the feed concentration, and V0 is the
feed volume. The benefits sought from higher concentrations can lead
to other problems such as reduced fluxes, larger area and pumps, pos-
sible denaturation, and extra lines that may have issues with cleaning
and product recovery. This not only has caused significant commis-
sioning and validation delays but also has led to the scrapping of a
process skid as unworkable. The number of pump passes will also be
higher, leading to greater potential protein degradation. Fed batch
and bypass should be used only when necessary.
Diafiltration After concentration has removed permeable com-
ponents from the system, a new buffer can be added to dilute the
retained product back to its original volume. Repeated application of
this procedure, known as batch diafiltration, is used to exchange one
buffer for another and is conveniently implemented at lab scale.
An operating mode known as constant volume diafiltration
involves adding a new buffer to the system while withdrawing per-
meate at the same rate. Figure 20-57 shows the relationship between
retentate concentration and diavolumes N = (buffer volumes
added)/(fixed retentate volume) for batch and constant volume diafil-
tration with S = 1 (Cheryan, 1998). For a fully passing solute (S = 1)
such as a buffer, the retentate concentration decays 10-fold with each
2.3 diavolumes. Partially retained solutes do not decay as quickly and
require more diavolumes to reach a final concentration target. A fully
retained solute (S = 0) maintains its retentate concentration constant
at the initial value. For fully passing solutes, > 4.6 diavolumes are
needed to achieve a specification of < 1 percent of the original buffer
MEMBRANE SEPARATION PROCESSES 20-43
FIG. 20-54 Stirred cell TFF module. (Courtesy Millipore Corporation.)
TABLE 20-18 Mass-Transfer Correlations
Geometry Flow a b c e Comments
Tube Laminar 1.62 0.33 0.33 0.33 Theoretical1
Slit Laminar 1.86 0.33 0.33 0.33 Theoretical1
Turbulent 0.023 0.80 0.33 — Theoretical2
Turbulent 0.023 0.875 0.25 — Theoretical3
Spacer — 0.664 0.50 0.33 0.50 Experimental4
Stirred cell Laminar 0.23 0.567 0.33 — Experimental5
Rotating Laminar 0.75 0.50 0.33 0.42 Experimental6
Leveque, Ann. Mines, 13, 201 (1928).
Gekas et al., J. Membr. Sci., 30, 153 (1987).
Deissler in Advances in Heat and Mass Transfer, Harnett (ed.), McGraw-Hill, New York, 1961.
DaCosta et al., J. Membr. Sci., 87, 79 (1994) with L = spacer mesh length/2, dh = 4e/[2/h + (1 − e)s] for spacer length channel
height h, porosity e, and specific area s. See Costa for additional correction kdc.
Smith et al., Chem. Eng. Prog. Symp. Ser., 64, 45 (1968).
Holeschovsky et al., AIChE J., 37, 1219 (1991).
20-44 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-55 Process configurations.
TABLE 20-19 Batch and Fed-Batch Performance Equations
and diafiltration Fed-batch concentration
Retentate concentration ci = ci0 X1 − Si
ci = ᎏ
ᎏ[1 − (1 − Si)e−Si(X − 1)
Retentate mass Mi = Mi0 e−Si(N+ln X)
Mi = ᎏ
ᎏ[1 − (1 − Si)e−Si(X − 1)
Sizing A = ᎏ
ᎏ + ᎏ
ᎏ A = ᎏ
Comments • Simplest system and control scheme • Required for 40 < X < 100
• Smaller area or process time • Minimize fed-batch ratio (= total
• Limited to X < 40 V0 / recycle tank V) to minimize area
• Can require larger area and time
• Integrate for J or use Cheryan approximation
• Test process performance using proposed
components. It is common to add an extra 1 to 2 diavolumes as a
safety factor to ensure complete buffer exchange. Note that incom-
plete mixing (due to dead legs and liquid droplets on tank walls)
becomes significant at high diavolumes (N > 10) and causes the
curves in Fig. 20-57 to flatten out.
System sizing involves integration of Eq. (20-69) using a flux model
to give Eq. (20-71), where Vp is the permeate volume and J is the aver-
age flux. Note the direct tradeoff between area and process time.
Table 20-19 shows the concentration and diafiltration steps separated
and processing time.
A = where = ͵
The formulas for concentration and diafiltration can be combined
for the entire process to derive an expression for the loss of product
in the permeate. This loss is shown in Fig. 20-58 and Table 20-19 for
different levels of processing and membrane retention characteristics
(R = 1 − S). Note that a membrane with 1 percent passage (99 percent
retention) can have yield losses much higher than 1 percent because
the protein is repeatedly cycled past the membrane during the entire
process with losses at every pass. A high-yielding process (< 1 percent
loss) requires membrane passage of < 0.1 percent (retention of > 99.9
The purification of a product p from an impurity i by an ultrafiltra-
tion process is shown in Fig. 20-59 and Eq. (20-72), where Ci0 and Cp0
are the initial concentrations (g/L) of the two components, Yp is the
yield of product in the retentate, and ψ = selectivity = Si /Sp, the ratio
of passages. High yields are obtained in purifying out small solutes
(high selectivity) but are compromised in removing larger impurities
with similar passages to the product.
Purification factor PF = =
ppm in product
ppm in feed
MEMBRANE SEPARATION PROCESSES 20-45
Tank Ratio = 0
Tank Ratio = 5
Tank Ratio = 10
Tank Ratio = 20
0 5 10 15 20 25
Overall Volumetric Concentration Factor, V0/(V0 − Vpermeate)
FIG. 20-56 Batch and fed-batch concentration.
Continuous Operation Continuous operation (also called feed-
and-bleed) involves the partial recycle of the retentate. The residence
time and number of pump passes are in between those of single-pass
and batch operation, depending on the fraction of retentate recycled.
Several continuous units can be plumbed with the retentate flow feed-
ing the next consecutive system. This configuration is commonly used
for large-scale membrane systems.
System Operation NFF system requirements depend on each
application but should generally allow for installation, flushing,
integrity testing, cleaning or sanitization, processing, recovery, and
Continued decline in performance indicates a membrane cleaning
or compatibility issue. The adequacy of the cleaning step is deter-
mined by the recovery of at least 80 percent of the initial normalized
water flux. Although some variability in water flux is typical, any con-
sistent decline reflects an inadequate cleaning procedure.
Process Development/Scale-up Feasibility is evaluated by using
small single-module systems. This involves selection of membrane
type, operating processing conditions, processing strategies (concen-
tration, diafiltration, etc.), cleaning procedures, performance evalua-
tion, and sensitivity to feedstock variations. Pilot operation employs a
larger skid, representative of planned commercial operation, running
representative feedstock over extended periods to assess longer-term
repeatable performance (product quality and retention, costs, yields,
process flux, and membrane integrity).
Module change-out is based on preset number of uses, time, or when
performance (retention, flux, integrity) drops below preset specifications.
Fluxes scale directly and volumes scale proportionately with the
feed volume. A safety factor will be built into the scale-up design so
that process times will not be the same.
REVERSE OSMOSIS (RO) AND NANOFILTRATION (NF)
RO and NF processes employ pressure driving forces of 0.3 to 10.5
MPa to drive liquid solvents (primarily water) through membranes
while retaining small solutes. Reverse-osmosis (RO) membranes are
commonly rated by a > 90 percent NaCl retention and retention of > 50
Da neutral organics while passing smaller liquid solvents. Nanofiltration
membranes (sometimes called low-pressure RO, loose RO, or hyperfil-
tration) have 20 to 80 percent NaCl retention and retain > 200 to 1000
Da neutral organics. Dissolved gases have low retention. Neutral or
undissociated solutes have lower retention than charged or dissociated
solutes. Although virus-retaining membranes are sometimes referred to
as nanofilters, in keeping with membrane scientist classification they are
described under ultrafiltration. RO membranes will remove bacteria
and viruses. Lab-scale sample preparation in dilute solutions can be run
as NFF in centrifuge tubes. TFF is used in large commercial processes
and small home and shipboard water purification systems.
Commercial interest in RO began with the first high-flux, high-
NaCl-retention Loeb-Sourirajan anisotropic cellulose acetate mem-
brane. Practical application began with the thin film composite (TFC)
membrane and implementation for seawater desalination at Jeddah,
Saudi Arabia [Muhurji et al., Desalination, 76, 75 (1975)].
Applications RO is primarily used for water purification: seawater
desalination (35,000 to 50,000 mg/L salt, 5.6 to 10.5 MPa operation),
brackish water treatment (5000 to 10,000 mg/L, 1.4 to 4.2 MPa oper-
ation), and low-pressure RO (LPRO) (500 mg/L, 0.3 to 1.4 MPa oper-
ation). A list of U.S. plants can be found at www2.hawaii.edu, and a
26 Ggal/yr desalination plant is under construction in Ashkelon,
Israel. Purified water product is recovered as permeate while the con-
centrated retentate is discarded as waste. Drinking water specifica-
tions of total dissolved solids (TDS) ≤ 500 mg/L are published by the
U.S. EPA and of ≤ 1500 mg/L by the WHO [Williams et al., chap. 24
in Membrane Handbook, Ho and Sirkar (eds.), Van Nostrand, New
York, 1992]. Application of RO to drinking water is summarized in
Eisenberg and Middlebrooks (Reverse Osmosis Treatment of Drink-
ing Water, Butterworth, Boston, 1986).
RO competes well with multistage flash evaporation in desalination
applications and is the current benchmark due to advantages in both
capital and operating costs. It may be too costly to compete against
freshwater sourced from wells and pipelines. LPRO and electrodialy-
sis compete for brackish water treatment applications. Compact home
kitchen RO runs on 0.4-MPa line pressure for < 2000-mg/L TDS feed
solutions to reduce salts, odor, and trace metals such as As. These
units range from $200 to $500 (U.S. 2005) to produce water at $0.06
to $0.25 per gallon. RO processes were estimated to represent 23 per-
cent of the 3 Ggal/day of worldwide potable water production in 1986
[Wangnick, IDA Magazine, 1, 53 (1986)].
Useful conversions are as follows:
Volume: 1 m3
= 1000 L = 264.2 U.S. gal = 8.11 × 10−4
acre·ft = 35.32 ft3
Pressure: 1 bar = 14.5 psi = 0.1 MPa
Flow: 1 U.S. gal/min = 1440 gal/day = 227 L/h = 5.47 m3
Concentration: 1 mg/L = 1 ppm = 17.1 grains/gal
Process water applications include boiler water feed pretreatment
before ion exchange or electrodialysis. RO is also used for ultrahigh-
purity water production for use in laboratories, medical devices (kid-
ney dialysis), microelectronic manufacturing (rinse fluids per ASTM
D-19 D5127-90, 1990), and pharmaceutical manufacturing (purified
water or water for injection as specified by USP).
RO applications recovering valuable product in the retentate
include tomato, citrus, and apple juice dewatering, sometimes in com-
bination with thermal evaporators at $0.15 to $0.18 per gallon [Short,
chap. 9 in Bioseparation Processes in Foods, Singh and Rezvi (eds.),
Marcel Dekker, New York, 1995]. Dealcoholization of wine or beer
involves diafiltering alcohol through RO membranes while retaining
flavor compounds in the retentate (Meier, Wine East, Nov./Dec., 10,
1991). LPRO is used to desalt soy sauce and dairy whey. Small-mole-
cule drugs (e.g., penicillin) can be recovered by RO membranes.
RO can recover metals, antifreeze, paint, dyes, and oils in the reten-
tate while generating cleaned up wastewater permeate for disposal.
RO is also used to reduce the volume of waste liquids (e.g., spent sul-
fite liquor in paper manufacturing). Wastewater treatment application
removals of 95 percent TOC, > 90 percent COD, > 98 percent PAH
compounds, and pesticides > 99 percent have been seen [Williams et
al., chap. 24 in Membrane Handbook, Sirkar and Ho (eds.), Van Nos-
Membranes, Modules, and Systems
Membranes RO membranes are designed for high salt reten-
tion, high permeability, mechanical robustness (to allow module fab-
rication and withstand operating conditions), chemical robustness (to
fabrication materials, process fluids, cleaners, and sanitizers), low
extractables, low fouling characteristics, high capacity, low cost, and
The predominant RO membranes used in water applications
include cellulose polymers, thin film composites (TFCs) consisting of
aromatic polyamides, and crosslinked polyetherurea. Cellulosic mem-
branes are formed by immersion casting of 30 to 40 percent polymer
lacquers on a web immersed in water. These lacquers include cellu-
lose acetate, triacetate, and acetate-butyrate. TFCs are formed by
interfacial polymerization that involves coating a microporous mem-
brane substrate with an aqueous prepolymer solution and immersing
in a water-immiscible solvent containing a reactant [Petersen, J. Membr.
Sci., 83, 81 (1993)]. The Dow FilmTec FT-30 membrane developed
by Cadotte uses 1-3 diaminobenzene prepolymer crosslinked with 1-3
and 1-4 benzenedicarboxylic acid chlorides. These membranes have
NaCl retention and water permeability claims.
Cellulose membranes can generally tolerate a pH of 3 to 6 and 0.3
to 1.0 ppm of chlorine while TFC membranes can generally tolerate a
pH of 3 to 11 and < 0.05 ppm of chlorine. Membrane tolerance is also
described as the permitted cumulative ppm-hours of membrane expo-
sure to chlorine. TFC membranes can range from 1000 to 12,000
ppm·h. Always check specific filter specifications. RO membranes
approved for food use by the FDA include aromatic polyamides,
polypiperazinamide, and aryl-alkyl polyetherurea (Code of Federal
Regulations, Title 21, Part 177, U.S. GPO, Washington, D.C., annually).
Nanofiltration membranes are negatively charged and reject multi-
valent anions at a much higher level than monovalent anions, an effect
described as Donnan exclusion. Nanofiltration membranes have
MgSO4 retention and water permeability claims.
Modules RO modules are available in spiral, hollow fiber, tubu-
lar, and plate and frame formats. A comparison of characteristics is
shown in Table 20-20. Spirals are used the most since they are com-
pact, easy-to-use modules; have low feed-side pressure drops; are less
prone to clogging; and are easily cleaned, mechanically robust, and
economical. The standard spiral module has an 8-in diameter, is 40 in
long with about 30-m3
area. An anti-telescoping end cap is used to
prevent the leaves from being pushed out of alignment by axial reten-
tate flow. Hollow fibers are run with the feed on the shell side, flow-
ing radially inward and out along the lumen. This places less
mechanical stress on the fibers. Tubular modules are used with high-
suspended-solids and high-viscosity fluids.
System Configurations Figure 20-60 shows common RO sys-
tem configurations for permeate product recovery. For single-pass
operation, the feed passes once through a membrane to generate
acceptable permeate product. Double-pass operation uses permeate
staging (the permeate from the first module feeds the second module)
with allowance for blending to generate acceptable product. This
configuration may be needed for high-concentration, hard-to-remove
impurities, high product purities, and lower-retention modules. Double-
pass operation has been recommended for desalination of > 38,000
mg/L [Klinko et al., Desalination, 54, 3 (1985)] or modules with salt
retention < 99.2 percent [Kaschemekat et al., Desalination, 46, 151
(1983)]. It is also common in ultrahigh-purity water systems.
Spirals are normally packed at 5 to 7 modules per cylindrical hous-
ing and operate with staged retentates where the retentate from the
first module feeds the second module. Retentate staging increases the
length of the feed channel. Retentate staging of housings is used to
increase the conversion ratio from around 25 to 50 percent in one
housing stage to 50 to 75 percent for two housing stages and 75 to 90
percent in three housing stages. Since the feed flow is reduced by the
MEMBRANE SEPARATION PROCESSES 20-47
90% 95% 100%
Product Yield in Retentate
=0 =1 =10 =500 =1000Ψ =
FIG. 20-59 Purification versus yield.
TABLE 20-20 RO Module Performance
Spiral Hollow fiber Tubular frame
Packing density, m2
800 6000 70 500
Feed flow, m3
⋅s) 0.3–0.5 0.005 1–5 0.3–0.5
Feed ∆P, psi 40–90 1.5–4.5 30–45 45–90
Plugging sensitivity High High Low Moderate
Cleanability Fair Poor Excellent Good
Relative cost Low Low High High
Chemistries available C and TFC C and TFC C only C and TFC
conversion ratio after passing through the first module stage, main-
taining a minimum crossflow rate requires that the second module
stage have less membrane area in what has been called a “Christmas
tree” design. Interstage pumping between housings may be required
to maintain pressures and flows.
Retentate product recovery employs the batch, fed-batch, or con-
tinuous processing systems, and concentration or diafiltration modes
of operation (see general membrane section). Beer and wine dealco-
holization uses batch diafiltration.
Component Transport in Membranes RO is analyzed as a
rate, not equilibrium, process. Polarization of retained salts and other
low-MW components is described in the general membrane section.
Solvent flux Jw and solute flux Ji through RO membranes can be
described by using Eq. (20-73), the solution diffusion imperfection
model [Sherwood et al., Ind. Eng. Chem. Fund., 6, 2 (1967)]. This
model considers transport through the membrane in parallel with
transport through a small number of nonretentive defects. Transport
of solvent or solute through the membrane occurs by partitioning
between the upstream polarization layer and membrane phases at the
membrane surface, diffusion through the membrane, and partitioning
between the membrane and downstream fluid. This process is
reflected in permeabilities Pw and P2. Dissolution in the membrane
and diffusion through it can be modeled as an activated process with
an associated inverse exponential temperature dependence of the
permeabilities. Transport of solvent or solute through defects occurs
by convective flow. Water viscosity increases roughly 2 percent/°C
Jw = (∆P − ∆π) + ∆P Ji = (cw − cp) + ∆Pcw (20-73)
In accordance with observed data, this model shows that water flux
Jw increases linearly with applied pressure ∆P, decreases with higher
salt concentration through its impact on osmotic pressure π, increases
with a smaller membrane thickness l, and increases with temperature
through the temperature dependence of the water permeability Pw.
The model also demonstrates that the solute or salt flux Ji increases lin-
early with applied pressure ∆P, increases with higher salt concentra-
tion cw, increases with a smaller membrane thickness l, and increases
with temperature through the temperature dependence of the solute
permeability P2. Polarization, as described early in this section, causes
the wall concentration cw to exceed the bulk concentration cb.
Osmotic Pressure The osmotic pressure π of salt solutions is cal-
π = − ln aw = ᎏ
for water activity aw, water partial molal volume V
w, gas constant R,
absolute temperature T, water molecular weight Mw, ion molality mi,
and osmotic coefficient φi for each ionic species i. Note that both Na+
are counted as separate species for NaCl. For seawater at
35,000 to 50,000 mg/L (ppm) and 25°C, the osmotic pressure is in the
range of 2.7 to 3.8 MPa (ϳ11 kPa/1000 ppm).
Solute Retention Retention is determined by the relative rates
of solute and solvent transport through the membrane. The impact of
operation on solute retention Ri = 1 − cp /cw can be evaluated from
Eq. (20-73) by using the mass balance Ji = cpJw. Assuming high reten-
tion, cp << cw, and the solvent flux through defects is negligible, so the
retention is calculated as
Ri = 1 − = 1 − (20-75)
At low pressures where ∆P ϳ ∆π, solvent flux is low but solute flux
is high (it is insensitive to pressure and occurs under a concentration
gradient). This results in low solute retention. Pressures of 1 MPa over
the feed osmotic pressure are needed to maintain high retention.
Increasing the applied pressure ∆P increases the solvent flux, diluting
the solute flux with solvent and resulting in a higher retention.
Increasing the feed solute concentration lowers the solvent flux
through its impact on osmotic pressure π, and decreases retention.
Increasing temperature increases the solvent permeability Pw, the
solute permeability P2, and the osmotic pressure. Solute retention
tends to drop slightly at higher temperature.
Ionic solutions require charge neutrality to prevent a higher flux of
species of one charge causing an associated electric current and volt-
age. This may limit the net salt flux through the membrane by the ion
with the lowest permeability so that the sodium ion will show different
retentions depending on the anion it is paired with. Charge neutrality
can also enrich the permeate concentration of a salt relative to the
feed to create a negative retention. For example, a feed solution con-
, and SO4
ions would show a permeate with
enriched levels of Cl−
to balance permeated Na+
Nanofilters incorporate negative membrane charge for higher
anion rejection. High feed salinities can passivate these charges and
reduce anion retention.
Pretreatment and Cleaning Pretreatment is commonly used to
extend membrane life and increase recovery. The representative pre-
treatment train for water purification applications in Fig. 20-61 con-
trols feed channel clogging, mineral scaling, fouling by organic films
and microorganisms, and oxidants that can degrade the membranes.
P2 + P3 ∆P
Pw(∆P − ∆π)
20-48 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-60 RO system configurations.
FIG. 20-61 RO pretreatment schematic.
Particle Removal Solid particles can plug feed channel spacer
channels and deposit as a cake layer on the membrane surface, caus-
ing increased flow resistance and lowering fluxes. Table 20-19 indi-
cates the sensitivity of different module types to feed channel
plugging. Particles may consist of organic colloids, rust flakes, precip-
itated iron hydroxide, algae, clay, colloidal silica, or silt. Pretreatment
includes using a combination of coarse screening, hydrocyclones, floc-
culent addition (e.g., alum, ferric sulfate, polyelectrolytes), a sand or
multimedia filtration bed with periodic backflushing, and depth filter
cartridges or ultrafiltration systems.
The ability of a feed solution to plug a membrane is determined by
the silt density index (SDI) employing a 0.45-µm PVDF microfilter at
30-psi differential pressure (ASTM Standard D-4189-82, 1987).
SDI = (20-76)
where Ti is the time it takes to filter the first 0.5 L and Tf is the time it
takes to filter 0.5 L starting when 15 min of test time has elapsed.
Table 20-21 is an interpretation of SDI values.
Antiscalants Precipitation of salts can form a scale on the mem-
brane surface when solubility limits (Table 20-22) are exceeded in the
retentate. This decreases flux.
Metal hydroxides (e.g., Fe, Mn, Al) can also be a problem (Rauten-
bach and Albrecht, Membrane Processes, Wiley, New York, 1989). A
chemical analysis of the feed solution composition along with consid-
eration of solubility products allows one to determine the significance
of precipitation. Solubility products can be affected by temperature,
pH, and ionic strength. Seasonal temperature variations must be con-
sidered. Concentrations of silica need to be < 120 mg/L in the feed.
Water hardness in milligrams per liter of total Ca2+
sidered soft at 0 to 17 mg/L, moderately hard at 60 to 120 mg/L, and
very hard at > 180 mg/L. The Langelier saturation index, a measure of
CaCO3 solubility, should be < 0 in the retentate to prevent precipita-
tion (Pure Water Handbook, Osmonics, 1997).
Scale prevention methods include operating at low conversion and
chemical pretreatment. Acid injection to convert CO3
to CO2 is com-
monly used, but cellulosic membranes require operation at pH 4 to 7
to prevent hydrolysis. Sulfuric acid is commonly used at a dosing of
0.24 mg/L while hydrochloric acid is to be avoided to minimize corro-
sion. Acid addition will precipitate aluminum hydroxide. Water soft-
ening upstream of the RO by using lime and sodium zeolites will
precipitate calcium and magnesium hydroxides and entrap some sil-
ica. Antiscalant compounds such as sodium hexametaphosphate,
EDTA, and polymers are also commonly added to encapsulate poten-
tial precipitants. Oxidant addition precipitates metal oxides for parti-
cle removal (converting soluble ferrous Fe2+
ions to insoluble ferric
Oxidant Removal The presence of oxidizers such as chlorine or
ozone can degrade polyamide RO membranes, causing a drop in salt
retention. Cellulosic membranes are less sensitive to attack. Addition
of 1.5 to 6 mg sodium bisulfite/ppm chlorine or contacting with acti-
vated carbon will remove oxidizers. Vacuum degassing with a hydropho-
bic filter module is also used.
Biocides Microorganisms can form biofilms on membrane sur-
faces, causing lower fluxes [Flemming in Reverse Osmosis, Z. Amjad
(ed.), Van Nostrand, New York, 1993, p. 163]. Cellulosic membranes
100(1 − TiրTf)
can be consumed by bacteria, causing lower retention. Control
involves pretreatment with chemicals or UV light and membrane
chemical sanitization on a regular or periodic basis. Chemical sani-
tants include chlorine, bromine, ozone and peroxide/peracetic acid
oxidants (subject to membrane limits), formaldehyde, sodium bisul-
fite, copper sulfate for algae and plankton control, or chloramines.
The use of oxidizers will also remove dissolved hydrogen sulfide. Cel-
lulosic membranes are commonly sanitized with 0.2 mg/L of free
Organic Foulants Organic components can range from 1 mg/L
in the open ocean to 80 to 100 mg/L in coastal waters. Components
include humic and fulvic acids (800 to 50,000 kDa) from decayed veg-
etation, or contaminating petroleum that can adsorb to the mem-
branes and reduce their flux. Carbon beds used for dechlorination can
also remove these organics by adsorption. For wastewater purification
to remove organics, high pH and ozonation improve retention and
reduce flux decline (Bhattacharyya et al., Separation of Hazardous
Organics by LPRO—Phase II, U.S. EPA, 1990).
Cleaning/Sanitization Membrane cleaning restores initial
membrane flux without compromising retention to extend membrane
life. Appropriate treatment depends on the nature of the foulants,
membrane compatibility, and compatibility of the chemicals with the
application. Cleaning generally uses acid flushing (for scalants) but
may include base or enzyme flushing (for organics) and 80°C hot
water flushing (for sanitization/microorganism control).
Design Considerations and Economics
Osmotic Pinch As the feed travels along the feed channels, fric-
tion losses lower the feed pressure. Solvent flux along the feed chan-
nel dewaters the feed and increases solute concentration and osmotic
pressure. For a recovery or conversion ratio CR, a module solute
retention R, and an inlet solute concentration of co, the bulk concen-
tration cb is
cb = co(1 − CR)−R
Both the feed pressure drop and the osmotic pressure increase
lower the driving force for solvent flow ∆P − ∆π along the feed chan-
nel, and the point at the end of the feed channel, having the lowest
driving force, is the osmotic pinch point. This point has both the low-
est solvent flux and the lowest solute retention. If the driving force
falls below 0 along the feed channel, the solvent flux will reverse (from
permeate into the feed) and subsequent membrane area will perform
“negative work.” Economic justification of membrane at the pinch
point requires that it generate adequate flux to pay for itself. This, in
turn, requires that the driving force be about 1 MPa.
Process Economics and Optimization Table 20-23 shows some
representative process economics for several large 1 Mgal/day system
designs [adapted from Ray, chap. 25 in Membrane Handbook, Ho and
Sirkar (eds.),Van Nostrand, New York, 1992]. These costs are a useful
guide, within ±25 percent, but specific quotes are needed to account
for site-specific differences between projects. No energy recovery
technologies are included. For capacities above 50 to 100 gal/day, sys-
tem costs scale almost linearly, as (gal/day capacity)0.85–0.95,
and can be
conveniently expressed as dollars per gallon per day. Total costs are
split evenly between capital and operating costs.
Energy use (pumping high flows to high pressures) represents the
biggest operating cost while membrane replacement operating costs
are only 10 to 15 percent. These operating costs can be represented as
$E ($/1000 gal) = (0.724 kWh/psig-1000 gal)
× (∆P psid)($/kWh)/CR-Ef (20-78)
where CR is the fractional conversion ratio (recovery), Ef is the pump
$A ($/1000 gal) = (0.432)($/m2
where J is the average system flux, L/(m2
⋅h), and T is the years of
By incorporating a flux model [Eq. (20-73)], solute concentration
[Eq. (20-77)], and polarization model [Eq. (20-60)] into the area cost
MEMBRANE SEPARATION PROCESSES 20-49
TABLE 20-21 SDI Interpretation
SDI value Operability
<1 Can run several years without colloidal fouling
1–3 Can run several months between cleaning
3–5 Frequent cleaning required
>5 Pretreatment required to reduce colloids
TABLE 20-22 Solubility Limits at 30°C
CaCO3 MgSO4 CaSO4 BaSO4 SiO2 SrSO4
65 mg/L 9 mg/L 2090 mg/L 2 mg/L 120 mg/L 114 mg/L
[Eq. (20-79)], the operating cost term ($A + $E) varies with ∆P and
CR. The osmotic pinch constraint requires a minimum ∆P to over-
come the average osmotic pressure and requires a maximum CR
where the concentrated retentate osmotic pressure exceeds the feed
pressure. Assuming that solute retention is sufficient to maintain per-
meate quality, there exist optimum conversion ratios and ∆P that
depend on the feed (osmotic pressure), module (mass transfer, per-
meability, life, cost), and pump (energy cost, efficiency). Conversion
ratios per module stage range from 25 to 50 percent, and typical inlet
pressures are shown in Table 20-23. No provision has been made here
for costs of disposing of the retentate.
One can also rearrange Eq. (20-79) to solve for ∆P in terms of area
and substitute in Eq. (20-78) to express the operating costs in terms of
membrane area. The energy operating cost is inverse in area, while
the membrane replacement cost is linear, reflecting the existence of
an optimum area. Addition of capital costs leads to total costs that
scale as b(area0.95
) + c(area−0.6
). Module vendors provide good general
recommendations and support to design and economic evaluation.
Recognizing that the high-pressure retentate contains considerable
kinetic energy before it is passed through a retentate valve provides an
opportunity for energy recycling. A variety of energy recovery technolo-
gies have been proposed including Pelton wheels, reverse-running
pumps, hydraulic turbochargers, and pressure exchangers with pump-
ing energy reductions of 22 to 60 percent [Darwish et al., Desalination,
75, 143 (1989)]. Energy recovery technologies are most useful for high-
pressure desalination but may not be worth the investment for brackish
water and LPRO applications (Glueckstern et al., chap. 12 in Synthetic
Membrane Processes, Academic Press, New York, 1984).
The capital cost of a 15 kgal/day RO skid, uninstalled with instru-
mentation, is about $35,000 (U.S., 2005) and scales with gallons-per-
day capacity to the 0.52 power [Ray, chap. 25 in Membrane Handbook,
Sirkar and Ho (eds.), Van Nostrand, New York, 1992]. Approximately
25 to 30 percent of this cost is associated with membranes and hous-
ings. Operating costs are roughly $0.30/1000 gal for energy and
membrane replacement, and $0.09/1000 gal for pretreatment, main-
tenance, and cleaning. High-purity water systems employing double-
pass staging operate at conversions of 70 to 80 percent in the first pass
and 70 to 75 percent in the second pass. Home units using line pres-
sure run at 5 to 15 percent conversion.
System Operation For corrosive pressurized seawater, 316L
stainless steel is commonly used for piping and housings although
specialized alloys may be needed (Oldfield et al., Desalination, 55,
261 (1985)]. Sensors are used to measure inlet pH and temperature
(to control scaling), pump pressure (to prevent surges and ensure
sufficient driving force), pump flow (to ensure crossflow to prevent
excessive polarization), permeate flow (to monitor performance), and
permeate conductivity (to ensure module integrity and product
quality). Control systems are used to monitor and display sensor val-
ues, track and optimize operation over time, trigger alarms, and store
the data. Provisions for module replacement and start-up are needed.
For batch processing, systems comparable to ultrafiltration are
Posttreatment of the permeate for potable water use can include dis-
solved CO2 removal to prevent corrosion (by aeration, lime treatment),
chlorination for microbial control, and oxygenation to improve taste.
Ultrafiltration processes (commonly UF or UF/DF) employ pressure
driving forces of 0.2 to 1.0 MPa to drive liquid solvents (primarily
water) and small solutes through membranes while retaining solutes
of 10 to 1000 Å diameter (roughly 300 to 1000 kDa). Commercial
operation is almost exclusively run as TFF with water treatment appli-
cations run as NFF. Virus-retaining filters are on the most open end of
UF and can be run as NFF or TFF. Small-scale sample preparation in
dilute solutions can be run as NFF in centrifuge tubes.
The first large commercial application of UF was paint recycling,
followed by dairy whey recovery in the mid-1970s. UF applications
are enabled by low-temperature and low-cost operation.
Applications Electrodeposition of cationic paint resin on auto-
mobiles (connected to the cathode) provides a uniform, defect-free
coating with high corrosion resistance, but carries with it about 50
percent excess paint that must be washed off. UF is used to maintain
the paint concentration in the paint bath while generating a permeate
that is used for washing. The spent wash is fed back into the paint path
(Zeman et al., Microfiltration and Ultrafiltration, Marcel Dekker,
New York, 1996).
During cheese making, the coagulated milk or curd is used to make
cheese while the supernatant whey is a waste product rich in salts,
proteins, and lactose. Whey concentration and desalting by UF pro-
duce a retentate product that can be used as an animal feed supple-
ment or food additive. The MMV process (Maubois et al., French
Patent 2,052,121) involves concentrating the milk by UF after cen-
trifugation to remove the cream and before coagulation to improve
yields and reduce disposal costs.
UF is used in biopharmaceutical purification (proteins, viral and
bacterial vaccines, nucleic acids) for initial concentration of clarified
fluids or lysates (to reduce subsequent column sizes and increasing
binding), to change buffer conditions (for loading on to columns), and
to concentrate and change buffers (for final formulation into a storage
buffer) [Lutz et al., in Process Scale Bioseparations for the Biophar-
maceutical Industry, Shukla, et al. (eds.), CRC Press, New York,
2006]. For these applications, the drug product is recovered in the
retentate. For drug products with potential viral contaminants
20-50 ALTERNATIVE SEPARATION PROCESSES
TABLE 20-23 Representative RO Process Costs
Costs Seawater Brackish water Softening Notes
Inlet pressure 6.9 MPa 3.1 MPa 1.4 MPa
Flux 25 LMH* 42 LMH 42 LMH
Conversion 40% 60% 75%
Total cost, $/1000 gal 4.7 2.3 1.3 15-yr life
Capital cost 2.1 1.1 0.6 12% interest
Operating cost 2.6 1.2 0.7 15% downtime
Total capital cost, $(gal/day) 4.5 2.3 1.3
Direct costs 3.7 1.9 1.1
Equipment 3.3 1.6 0.8
Indirect costs 0.8 0.4 0.2
Total operating cost, $/1000 gal 2.6 1.2 0.7
Energy 1.6 0.4 0.2 $0.075/kWh
Membrane replacement 0.4 0.2 0.2 3-yr life
Chemicals 0.2 0.2 0.1 2 per shift @$14/h
Labor 0.3 0.3 0.2
Other 0.1 0.1 0.0
*LMH = liters per square meter per hour.
(sourced from mammalian cells or human and animal blood plasma),
virus removal filters employed in either a TFF or NFF mode retain
the viruses while the drug product is recovered in the permeate.
UF is used to clarify various fruit juices (apple, grape, pear, pineap-
ple, cranberry, orange, lemon) which are recovered as the permeate
[Blanck et al., AIChE Symp. Ser. 82, 59 (1986)]. UF has also been
used to remove pigments and reducing browning in wine production
[Kosikowski in Membrane Separations in Biotechnology, McGregor
(ed.), Marcel Dekker, New York, 1986].
Wastewater treatment and water purification applications employ
UF in a TFF or NFF mode to produce permeate product with
reduced colloids, pyrogens, and viruses. Oil droplets in wastewater are
retained by UF for recycle or disposal at a significantly reduced vol-
Membranes, Modules, and Systems
Membranes UF membranes consist primarily of polymeric
structures (polyethersulfone, regenerated cellulose, polysulfone,
polyamide, polyacrylonitrile, or various fluoropolymers) formed by
immersion casting on a web or as a composite on a MF membrane.
Hydrophobic polymers are surface-modified to render them hydrophilic
and thereby reduce fouling, reduce product losses, and increase flux
[Cabasso in Ultrafiltration Membranes and Applications, Cooper
(ed.), Plenum Press, New York, 1980]. Some inorganic UF mem-
branes (alumina, glass, zirconia) are available but only find use in cor-
rosive applications due to their high cost.
Early ultrafiltration membranes had thin surface retentive layers
with an open structure underneath, as shown in Fig. 20-62. These
membranes were prone to defects and showed poor retention and
consistency. In part, retention by these membranes would rely on
large retained components in the feed that polarize or form a cake
layer that plugs defects. Composite membranes have a thin retentive
layer cast on top of a microfiltration membrane in one piece. These
composites demonstrate consistently high retention and can be
integrity-tested by using air diffusion in water.
Table 20-24 compares properties of commonly used polyethersul-
fone (PES) or regenerated cellulose membranes. Membrane selection
is based on experience with vendors, molecular weight rating for high
yields, chemical and mechanical robustness during product processing
and clean-in-place (extractables, adsorption, swelling, shedding, class
VI), process flux for sizing and costing, and the quality/consistency
(ISO, cGMP) of the vendor and the membrane. Membranes are
described by their normal water permeability (NWP), the ratio of their
flux at 25°C to the average transmembrane pressure (TMP). Measure-
ments taken at different temperatures can be adjusted to 25°C using a
viscosity ratio correction. Process fluxes will be somewhat lower due to
fouling. Membranes containing a hydrophilic surface generally have
lower fouling and produce higher process fluxes.
UF membranes are assigned a nominal molecular weight limit
(NMWL) corresponding roughly to their ability to retain a solute of a
particular molecular weight. While tight membranes (i.e., low-
molecular-weight cutoff ratings) provide high retention, they have
corresponding low flow rates requiring greater membrane area to do
the job. This leads to larger pumps with large holdup volumes and
potential negative impact on product quality. A rule of thumb for
selecting membrane NMWL is to take 0.2 to 0.3 of the MW of the
solute that is to be retained by the membrane. However, the NMWL
values should only be used as a very rough guide or screening method
to select membranes for testing. Retention is based on hydrodynamic
size, not molecular weight, so linear chain dextrans show a higher pas-
sage than globular proteins of the same molecular weight. In addition,
there is no standardization of rating methods so specific retention
properties vary considerably among membranes and vendors depend-
ing on the marker solute selected (e.g., protein, dextran in a particular
buffer) and the level of retention selected for the marker solute (90,
95, 99 percent). Retention is affected by fouling due to adsorbed
components and polarized solutes on the membrane surface. See
retention below for additional discussion.
Modules and Systems UF modules include cassettes, spirals,
hollow fibers, tubes, flat sheets, and inorganic monoliths. These are
primarily run in TFF operation to increase flux by reducing plugging.
Exceptions run in NFF operation include virus removal and water
treatment applications where both cartridges and hollow fibers are
used. Continuous operation provides lower-cost operation, but the
biologicals process requires batch processing to meet regulatory
requirements. Applications with high solids require multiple-pass
operation to obtain a significant conversion ratio, but water treatment
MEMBRANE SEPARATION PROCESSES 20-51
FIG. 20-62 Ultrafiltration membranes. (Courtesy Millipore Corporation.)
TABLE 20-24 Ultrafiltration Membrane Properties
Material Advantages Disadvantages
Polyethersulfone (PES) Resistance to temperature, Cl2, pH, Hydrophobic
Regenerated cellulose Hydrophilic, low-fouling Sensitive to temperature, pH, Cl2, microbial
attack, mechanical creep
Polyamide Sensitive to Cl2, microbial attack
Inorganic Resistance to temperature, Cl2, pH,
high pressure, long life Cost, brittleness, high crossflow rates
can achieve high conversion in one pass. The properties of these mod-
ules and system operation are described in the general section. Table
20-25 indicates where these modules are used.
Component Transport in Membranes UF is analyzed as a
rate, not equilibrium, process. Pure solvent transport in UF is charac-
terized by J = TMPրRmembrane, where TMP (transmembrane pressure)
is the pressure difference across the membrane and R is the mem-
brane resistance. Resistance R can be related to membrane properties
by using Eq. (20-64) with pore size distributions following a lognormal
distribution [Zydney et al., J. Membr. Sci., 91, 293 (1994)]. At ionic
strengths < 0.1 M, charged membranes generate a streaming potential
across their thickness which can lead to a 15 percent or more dimin-
ished solvent flux by counterelectroosmosis (Newman, Electrochemi-
cal Systems, Prentice-Hall, Englewood Cliffs, N.J., 1973). Most UF
process fluids are aqueous with higher ionic strengths, but permeabil-
ity measurements using low-conductivity water can encounter this
Figure 20-63 shows that in the presence of retained solutes, the sol-
vent flux flattens out at high pressures. UF has a high level of polar-
ization where the membrane surface concentration can range up to
100 times the bulk concentration. Retained antibodies at a wall con-
centration of 191 g/L have an osmotic pressure Π of 30 psi (Mitra,
1978). That is, an elevated pressure of 30 psig must be applied to the
protein-rich retentate side of a water-permeable membrane contain-
ing 191 g/L of antibody in order to prevent water backflow from the
permeate side of the membrane containing water at 0 psig. This
diminishes the driving force for flow and leads to the mechanistic-
based osmotic flux model (Vilker, 1984):
Polarization flux model J = (20-80)
where R is the intrinsic membrane retention and ∆∏(Cw) is the
osmotic pressure at the wall concentration. An empirically based flux
model can also be defined by omitting the osmotic term and adding a
compressible polarization resistance term in the denominator
TMP − R ⋅ ∆Π(cw)
µ ⋅ (Rmembrane + Rfouling)
Equation (20-80) requires a mass transfer coefficient k to calculate
cw and a relation between protein concentration and osmotic pressure.
Pure water flux obtained from a plot of flux versus pressure is used to
calculate membrane resistance (typically small). The LMH/psi slope is
referred to as the NWP (normal water permeability). The membrane
plus fouling resistances are determined after removing the reversible
polarization layer through a buffer flush. To illustrate the components
of the osmotic flux model, Fig. 20-63 shows flux versus TMP curves
corresponding to just the membrane in buffer (Rfouling = 0, cw = 0),
fouled membrane in buffer (cw = 0), and fouled membrane with
The region at low flux/low TMP is called the linear region and is
dependent on TMP but independent of crossflow and bulk concen-
tration. The region at high flux/high TMP is called the polarized
region and is independent of TMP but dependent on crossflow and
bulk concentration. This extremely counterintuitive result is the con-
sequence of polarization. In between these two regions lies what is
termed the “knee” of the flux curve. Increasing pressure beyond the
knee gives diminishing returns in improving flux.
In the polarized region, the empirical gel model of Eq. (20-81) can
provide a good description of the flux behavior. The gel model is based
on polarization [Eq. (20-58)] with a constant wall concentration Cg
and complete retention. The mass-transfer coefficient k is dependent
on crossflow velocity, as shown in Table 20-18 for different geometries
and flow regimes. Concentration-dependent transport (viscosity and
diffusivity) will alter the velocity exponents. Operating at the knee of
the flux curve, either by design or due to fouling that causes a shift in
the flux curve to the right, will decrease the velocity dependence.
Note that the empirical gel point concentration does not correspond
to an actual gel on the membrane surface. As the feed concentration
rises to within 5 to 10 percent of the gel point, the flux tends to drop
off less rapidly than Eq. (20-81) would predict. The gel point is insen-
sitive to crossflow but can increase with temperature. Cheryan sum-
marizes gel point and velocity exponent values for a variety of
J = k ln cg րc (20-81)
20-52 ALTERNATIVE SEPARATION PROCESSES
TABLE 20-25 UF Modules and Systems
Application Paint recovery Whey processing Biologicals Water treatment Juice processing
Membrane PES PES Cellulose, PES, PVDF PES PES, PVDF
Module Fibers Spirals Cassette Fibers Tubes
System Continuous Continuous Batch, NFF Continuous, single-pass Batch
Characteristics Cost, plugging Cost Recovery Cost Plugging, cost
800 20 40 60
FIG. 20-63 Ultrafiltration flux behavior.
Solute Flux Solute partitioning between the upstream polariza-
tion layer and the solvent-filled membrane pores can be modeled by
considering a spherical solute and a cylindrical pore. The equilibrium
partition coefficient φ (pore/bulk concentration ratio) for steric exclu-
sion (no long-range ionic or other interactions) can be written as
Solute partitioning φ = (1 − λ)2
where λ = solute radius/pore radius.
Solute flux within a pore can be modeled as the sum of hindered
convection and hindered diffusion [Deen, AIChE J., 33, 1409 (1987)].
Diffusive transport is seen in dialysis and system start-up but is negli-
gible for commercially practical operation. The steady-state solute
convective flux in the pore is Js = KcJc = φKcJcw, where c is the radially
averaged solute concentration and Kc is the convective hindrance fac-
tor (from pore wall drag, solute lag, and long-range solute pore wall
and solute-solute potential interactions). For uncharged cylindrical
pores φKc = φ(2 − φ)exp(−0.7146λ2
) to within 2 percent [Zeman et al.,
in Synthetic Membranes, vol. 2, Turbak (ed.), ACS Symp. Ser., 54,
ACS, Washington, D.C. (1981), p. 412].
A charged solute in a pore has a higher potential energy due to the
distortion of its electrical field when it penetrates the pore wall into
the polymer. This effect increases solute retention. The apparent size
of solutes can be magnified by more than an order of magnitude by
this effect. See Deen for other pore geometries and the impacts of
long-range solute pore wall and solute-solute interactions. The use of
electrostatic interactions to separate solutes is discussed by van Reis
[Biotechn. and Bioeng., 56(1), 71 (1997)].
Solute Passage/Retention The intrinsic passage (cp/cw) is deter-
mined by the ratio of the solute flux to the solvent flux as si = φKc. The
intrinsic passage is inherent to the membrane and solute. As shown in
Fig. 20-64 for a single pore, solute passage (or retention) does not fol-
low a step change. Figure 20-65 shows passage for a membrane with a
pore size distribution [Tkacik et al., Biotechnology, 9, 941 (1991)].
The observed passage (cp/cb) varies with both the intrinsic passage
and the extent of polarization, as shown in Eq. (20-83):
Observed passage so = (20-83)
1/si − 1
⋅ exp (− Jրk)
At low flux these passages are equivalent. At high fluxes the wall
concentration is high due to polarization and the observed passage
increases, approaching 100 percent, regardless of the intrinsic
In multicomponent systems, large solutes with lower diffusivity,
polarize more and exclude smaller solutes from the membrane sur-
face, decreasing their passage. Operation at the “knee” of the flex
curve reduces this effect.
Design Considerations and Economics The selection among
batch, fed-batch, single-pass, or continuous operation has been
described earlier in this section. Important performance considera-
tions for UF applications involve product purity and/or concentration,
product yield, and cost/sizing for a given production rate.
Batch or Fed-Batch Operation This mode of operation is typi-
cal of biologicals and juice processing where high solids and low fluxes
require multiple passes, and batch operation is characteristic of the
manufacturing process. Formulas in Table 20-19 can be used to cal-
culate the required volume reduction factor X, diafiltration volumes
MEMBRANE SEPARATION PROCESSES 20-53
0.00 0.20 0.40 0.60 0.80 1.00
Particle diameter/pore diameter
Spherical particle passage
in a cylindrical pore
FIG. 20-64 Solute passage.
30K Composite Regenerated Cellulose (PLCTK)
Dextran Retention Challenge
1,000 10,000 100,000
FIG. 20-65 Intrinsic membrane retention.
N, and solute passage Si needed to produce desired retentate product
with impurity concentrations ci and retentate product yield Mi/Mi0.
Permeate product characteristics for batch operation can be deter-
mined by mass balances using a permeate volume of Vp = V0(1 −
1րX + NրX), a mass of solute i in the permeate as Mi, permeate =
Mi0(1 − e−Si(N+lnX)
) and the permeate concentration as the ratio of the
mass to volume. Fed-batch equations are obtained by using the same
The system area A needed for batch or fed-batch operation can be
calculated by using the formulas in Table 20-19 for production rates
V0/t based on feed volume V0 and average fluxes during process steps.
The final retentate batch volume VR = V0րX or permeate batch volume
Vp = V0(1 − 1րX + NրX) can be used to restate the production rate on
other bases. Although experience can be used to estimate solute pas-
sage and process fluxes, they should be determined experimentally for
Single-Pass TFF Single-pass TFF or NFF operation is typical of
water treatment. Low solids allow high fluxes with high permeate
product recovery and low retentate disposal costs. NFF operation is
also used for virus retention in small-batch biologicals processing
where low NFF fluxes (high areas) are compensated by ease of use,
low residence time, and low capital costs.
Equations (20-66) and (20-67) present single-pass formulas relating
retentate solute concentration, retentate crossflow, permeate flow,
and membrane area. For relevant low-feed-concentration applica-
tions, polarization is minimal and the flux is mainly a function of pres-
sure. Spiral or hollow fiber modules with low feed channel and
permeate pressure drops are preferred.
Continuous Operation This mode of operation is typical of paint
recovery, whey processing, and wastewater processing where high
solids and low fluxes require multiple passes, and continuous operation
is allowed by the manufacturing process. For a feed concentration cF, a
volumetric concentration factor X (= retentate flow/feed flow), and
retention R, the outlet retentate concentration is
c = cF
΄ ΅ (20-84)
In comparison to batch processing, this process operates at the final
concentration with associated low fluxes. The use of multiple stages
where the retentate outlet from one stage feeds the next allows suc-
cessive systems to operate at different fluxes and will improve mem-
Process Description Microfiltration (MF) separates particles
from true solutions, be they liquid or gas phase. Alone among the
membrane processes, microfiltration may be accomplished without
the use of a membrane. The usual materials retained by a microfiltra-
tion membrane range in size from several µm down to 0.2 µm. At the
low end of this spectrum, very large soluble macromolecules are
retained by a microfilter. Bacteria and other microorganisms are a
particularly important class of particles retained by MF membranes.
Among membrane processes, dead-end filtration is uniquely common
to MF, but cross-flow configurations are often used.
Brief Examples Microfiltration is the oldest and largest mem-
brane field. It was important economically when other disciplines
were struggling for acceptance, yet because of its incredible diversity
and lack of large applications, it is the most difficult to categorize.
Nonetheless, it has had greater membrane sales than all other mem-
brane applications combined throughout most of its history. The early
success of microfiltration was linked to an ability to separate microor-
ganisms from water, both as a way to detect their presence, and as a
means to remove them. Both of these applications remain important.
Laboratory Microfiltration membranes have countless laboratory
uses, such as recovering biomass, measuring particulates in water, clari-
fying and sterilizing protein solutions, and so on. There are countless
examples for both general chemistry and biology, especially for analytical
procedures. Most of these applications are run in dead-end flow, with the
membrane replacing a more conventional medium such as filter paper.
X − R(X − 1)
Medical MF membranes provide a convenient, reliable means to
sterilize fluids without heat. Membranes are used to filter injectable
fluids during manufacture. Sometimes they are inserted into the tube
leading to a patient’s vein.
Process Membrane microfiltration competes with conventional
filtration, particularly with diatomaceous earth filtration in general-
process applications. A significant advantage for membrane MF is the
absence of a diatomaceous earth residue for disposal. Membranes
have captured most of the final filtration of wine (displacing asbestos),
are gaining market share in the filtration of gelatin and corn syrup
(displacing diatomaceous earth), are employed for some of the cold
pasteurization of beer, and have begun to be used in the pasteuriza-
tion of milk. Wine and beer filtration operate dead-end; gelatin, corn
syrup, and milk are cross-flow operations. MF is used to filter all fluid
reactants in the manufacture of microcircuits to ensure the absence of
particulates, with point-of-use filters particularly common.
Gas Phase Microfiltration plays an important and unique role in
filtering gases and vapors. One important example is maintaining
sterility in tank vents, where incoming air passes through a microfilter
tight enough to retain any microorganisms, spores, or viruses. A
related application is the containment of biological activity in purge
gases from fermentation. An unrelated application is the filtration of
gases, even highly reactive ones, in microelectronics fabrication to
prevent particulates from contaminating a chip.
Downstream Processing Microfiltration plays a significant role
in downstream processing of fermentation products in the pharma-
ceutical and bioprocessing industry. Examples are clarification of
fermentation broths, sterile filtration, cell recycle in continuous fer-
mentation, harvesting mammalian cells, cell washing, mycelia recov-
ery, lysate recovery, enzyme purification, vaccines, and so forth.
MF Membranes Microfiltration is a mature field that has prolif-
erated and subdivided. The scope and variety of MF membranes far
exceeds that in any other field. A good overview is given by Strath-
mann [in Porter (ed.), op. cit., pp. 1–78]. MF membranes may be clas-
sified into those with tortuous pores or those with capillary pores.
Tortuous-pore membranes are far more common, and are spongelike
structures. The pore openings in MF are much larger than those in
any other membrane. Surface pores may be observed by electron
microscope, but tortuous pores are much more difficult to observe
directly. Membranes may be tested by bubble-point techniques.
Many materials not yet useful for tighter membranes are made into
excellent MF membranes. Retention is the primary attribute of an
MF membrane, but important as well are permeability, chemical and
temperature resistance, dirt capacity (for dead-end filters), FDA-USP
approval, inherent strength, adsorption properties, wetting behavior,
and service life.
Membrane-production techniques listed below are applicable pri-
marily or only to MF membranes. In addition, the Loeb-Sourirjain
process, used extensively for reverse osmosis and ultrafiltration mem-
branes, is used for some MF membranes.
Membranes from Solids Membranes may be made from micro-
particles by sintering or agglomeration. The pores are formed from
the interstices between the solid particles. The simplest of this class
of membrane is formed by sintering metal, metal oxide, graphite,
ceramic, or polymer. Silver, tungsten, stainless steel, glass, several
ceramics, and other materials are made into commercial membranes.
Sintered metal may be coated by TiO2 or zirconium oxide to produce
MF and UF membranes. Membranes may be made by the careful
winding of microfibers or wires.
Ceramic Ceramic membranes are made generally by the sol-gel
process, the successive deposition of ever smaller ceramic precursor
spheres, followed by firing to form multitube monoliths. The diame-
ter of the individual channels is commonly about 2 to 6 mm. Mono-
liths come in a variety of shapes and sizes. A 19-channel design is
common. One manufacturer makes large monoliths with square
Track-Etched Track-etched membranes (Fig. 20-66) are now
made by exposing a thin polymer film to a collimated beam of radia-
tion strong enough to break chemical bonds in the polymer chains.
The film is then etched in a bath which selectively attacks the dam-
aged polymer. The technique produces a film with photogenic pores,
20-54 ALTERNATIVE SEPARATION PROCESSES
whose diameter may be varied by the intensity of the etching step.
Commercially available membranes have a narrow pore size distribu-
tion and are reportedly resistant to plugging. The membranes have
low flux, because it is impossible to achieve high pore density without
sacrificing uniformity of diameter.
Chemical Phase Inversion Symmetrical phase-inversion mem-
branes (Fig. 20-67) remain the most important commercial MF mem-
branes produced. The process produces tortuous-flow membranes. It
involves preparing a concentrated solution of a polymer in a solvent.
The solution is spread into a thin film, then precipitated through the
slow addition of a nonsolvent, usually water, sometimes from the
vapor phase. The technique is impressively versatile, capable of pro-
ducing fairly uniform membranes whose pore size may be varied
within broad limits.
Thermal Phase Inversion Thermal phase inversion is a tech-
nique which may be used to produce large quantities of MF mem-
brane economically. A solution of polymer in poor solvent is prepared
at an elevated temperature. After being formed into its final shape, a
sudden drop in solution temperature causes the polymer to precipi-
tate. The solvent is then washed out. Membranes may be spun at high
rates using this technique.
Stretched Polymers MF membranes may be made by stretching
(Fig. 20-68). Semicrystalline polymers, if stretched perpendicular to
the axis of crystallite orientation, may fracture in such a way as to make
reproducible microchannels. Best known are Goretex® produced
from Teflon®, and Celgard® produced from polyolefin. Stretched
polymers have unusually large fractions of open space, giving them
very high fluxes in the microfiltration of gases, for example. Most such
materials are very hydrophobic.
Membrane Characterization MF membranes are rated by flux
and pore size. Microfiltration membranes are uniquely testable by
direct examination, but since the number of pores that may be
observed directly by microscope is so small, microscopic pore size
determination is mainly useful for membrane research and verifica-
tion of other pore-size-determining methods. Furthermore, the most
critical dimension may not be observable from the surface. Few MF
membranes have neat, cylindrical pores. Indirect means of measure-
ment are generally superior. Accurate characterization of MF mem-
branes is a continuing research topic for which interested parties
should consult the current literature.
Bubble Point Large areas of microfiltration membrane can be
tested and verified by a bubble test. Pores of the membrane are
filled with liquid, then a gas is forced against the face of the mem-
brane. The Young-Laplace equation, ∆P = (4γ cos Θ)/d, relates the
pressure required to force a bubble through a pore to its radius, and
the interfacial surface tension between the penetrating gas and the
liquid in the membrane pore. γ is the surface tension (N/m), d is the
pore diameter (m), and P is transmembrane pressure (Pa). Θ is the
liquid-solid contact angle. For a fluid wetting the membrane per-
fectly, cos Θ = 1.
By raising the gas pressure on a wet membrane until the first bub-
ble appears, the largest pore may be identified, and its size computed.
This is a good test to run on a membrane apparatus used to sterilize a
fluid, since bacteria larger than the identified largest pore (or leak)
cannot readily penetrate the assembly. Pore-size distribution may also
be run by bubble point. Bubble-point testing is particularly useful in
assembled microfilters, since the membrane and all seals may be ver-
ified. Periodic testing ensures that the assembly retains its integrity.
Diffusional flow of gas is a complication in large MF assemblies. It
results from gas dissolving in pore liquid at the high-pressure side, and
desorbing at the low-pressure side. If the number of pores and the
average pore length are known, the effect can be computed. Special
protocols are used when this method is used for critical applications.
Detail is provided in ASTM F316-86, “Standard test method for pore
size characteristics of membrane filters by bubble point and mean
flow pore test.” The bubble-point test may also be run using two liq-
uids. Because interfacial surface tensions of liquids can be quite low,
this technique permits measurements on pores as small as 10 nm.
Charged Membranes The use of tortuous-flow membranes con-
taining a positive electrical charge may reduce the quantity of nega-
tively charged particles passing even when the pore size is much larger
than the particle. The technique is useful for making prefilters or lay-
ered membranes that withstand much higher solids loadings before
Bacteria Challenge Membranes are further tested by chal-
lenge with microorganisms of known size: their ability to retain all of
the organisms is taken as proof that all pores are smaller than the
organism. The best-known microorganism for pore-size determina-
tion is Pseudomonas diminuta, an asporogenous gram-negative rod
with a mean diameter of 0.3 µm. Membranes with pore size smaller
than that are used to ensure sterility in many applications. Leahy and
MEMBRANE SEPARATION PROCESSES 20-55
FIG. 20-66 Track-etched 0.4-µm polycarbonate membrane. (Courtesy Milli-
FIG. 20-67 Chemical phase inversion 0.45-µm polyvinylidene fluoride mem-
brane. (Courtesy Millipore Corporation.)
FIG. 20-68 Stretched polytetrafluoroethylene membrane. (Courtesy Milli-
Sullivan [Pharmaceutical Technology, 2(11), 65 (1978)] provide
details of this validation procedure.
Membrane thickness is a factor in microbial retention. Tortuous-
pore membranes rated at 0.22 µm typically have surface openings as
large as 1 µm (Fig. 20-67). Narrower restrictions are found beneath
the surface. In challenge tests, P. diminuta organisms are found well
beneath the surface of a 0.2-µm membrane, but not in the permeate.
Latex Latex particles of known size are available as standards.
They are useful to challenge MF membranes.
Process Configuration As befits a field with a vast number of
important applications and a history of innovation, there are countless
variations on how an MF process is run.
Dead-end versus Cross-flow Conventional filtration is usually
run dead-end, and is facilitated by amendments that capture the
particulates being removed. Membranes have very low dirt capacity, so
only applications with very low solids to be removed are run in
conventional dead-end flow. A rough upper limit to solids content is
about 0.5 percent; streams containing <0.1 percent are almost always
processed by dead-end devices. Since dead-end membrane equipment is
much less expensive than cross-flow, great ingenuity is applied to pro-
tecting the critical membrane pores by structured prefilters to remove
brane may be run inverted. A review of dead-end membrane filtration is
given by Davis and Grant [in Ho and Sirkar (eds.), op. cit., pp. 461–479].
Cross-flow is the usual case where cake compressibility is a prob-
lem. Cross-flow microfiltration is much the same as cross-flow ultra-
filtration in principle. In practice, the devices are often different.
As with UF, spiral-wound membranes provide the most economical
configuration for many large-scale installations. However, capillary
devices and cassettes are widely employed, especially at smaller scale.
A detailed description of cross-flow microfiltration had been given by
Murkes and Carlsson [Crossflow Filtration, Wiley, New York (1988)].
Membrane Inverted Most membranes have larger openings on
one face than on the other. Common practice is to run the tightest
face against the feed in order to avoid plugging of the backing by par-
ticles. The rationale is that anything that makes it past the “skin” will
have relatively unimpeded passage into the backing and out with the
permeate. For very low solids this convention is reversed, the ratio-
nale being that the porous backing provides a trap for particulates,
rather like filter aid.
If the complete passage of soluble macromolecules is required, a
highly polarized membrane is an advantage. The upside-down mem-
brane hinders the back diffusion of macrosolutes. Countering the ten-
dency of the retained particulates to “autofilter” soluble macrosolutes,
the inhibition of back diffusion raises the polarization and thus the
passage of macrosolutes such as proteins. The particulates are physi-
cally retained by the membrane. Blinding and plugging can be con-
trolled if the membrane is backwashed frequently. This technique has
been demonstrated at high solids loadings in an application where
high passage of soluble material is critical, the microfiltration of beer
[Wenten, Rasmussen, and Jonsson, North Am. Membrane Soc. Sixth
Annual Meeting, Breckenridge, CO (1994)].
Liquid Backpulse Solid membranes are backwashed by forcing
permeate backward through the membrane. Frequent pulsing seems
to be the key.
Air Backflush A configuration unique to microfiltration feeds
the process stream on the shell side of a capillary module with the per-
meate exiting the tube side. The device is run as an intermittent dead-
end filter. Every few minutes, the permeate side is pressurized with
air. First displacing the liquid permeate, a blast of air pushed back-
ward through the membrane pushes off the layer of accumulated
solids. The membrane skin contacts the process stream, and while
being backwashed, the air simultaneously expands the capillary and
membrane pores slightly. This momentary expansion facilitates the
removal of imbedded particles.
Process Limitations The same sorts of process limitations
affecting UF apply to MF. The following section will concentrate on
Concentration Polarization The equations governing cross-flow
mass transfer are developed in the section describing ultrafiltration.
The velocity, viscosity, density, and channel-height values are all similar
to UF, but the diffusivity of large particles (MF) is orders-of-magnitude
lower than the diffusivity of macromolecules (UF). It is thus quite sur-
prising to find the fluxes of cross-flow MF processes to be similar to, and
often higher than, UF fluxes. Two primary theories for the enhanced
diffusion of particles in a shear field, the inertial-lift theory and the
shear-induced theory, are explained by Davis [in Ho and Sirkar (eds.),
op. cit., pp. 480–505], and Belfort, Davis, and Zydney [J. Membrane.
Sci., 96, 1–58 (1994)]. While not clear-cut, shear-induced diffusion is
quite large compared to Brownian diffusion except for those cases with
very small particles or very low cross-flow velocity. The enhancement of
mass transfer in turbulent-flow microfiltration, a major effect, remains
Fouling Fouling affects MF as it affects all membrane processes.
One difference is that the fouling effect caused by deposition of a
foulant in the pores or on the surface of the membrane can be con-
founded by a rearrangement or compression of the solids cake which
may form on the membrane surface. Also, the high, open space found
in tortuous-pore membranes makes them slower to foul and harder to
Equipment Configuration Since the early days when mem-
brane was available only in flat-sheet form, the variety of offerings of
various geometry and fabricated filter component types has grown
geometrically. An entire catalog is devoted just to list the devices
incorporating membranes whose area ranges from less than 1 cm2
to 3 m2
. Microfiltration has grown to maturity selling these relatively
small devices. Replacement rather than reuse has long been the cus-
tom in MF, and only with later growth of very large applications, such
as water, sewage, and corn sweeteners, has long membrane life
become an economic necessity on a large scale.
Conventional Designs Designs familiar from other unit opera-
tions are also used in microfiltration. Cartridge-filter housings may be
fitted with pleated MF membrane making a high-area dead-end
membrane filter. Plate-and-frame type devices are furnished with MF
sheet stock, and are common in some applications. Capillary bundles
with tube-side feed are used for cross-flow applications, and are occa-
sionally used in dead-end flow. A few tubular membranes are used.
Spiral-wound modules are becoming increasingly important for process
applications where economics are paramount. Belt filters have been
made using MF membrane.
Ceramics Ceramic microfilters for commercial applications
are almost always employed as tube-side feed multitube monoliths.
They are also available as flat sheet, single tubes, discs, and other
forms primarily suited to lab use. They are used for a few high-
temperature applications, in contact with solvents, and particularly at
very high pH.
Cassettes Cassette is a term used to describe two different cross-
flow membrane devices. The less-common design is a usually large
stack of membrane separated by a spacer, with flow moving in parallel
across the membrane sheets. This variant is sometimes referred to as a
flat spiral, since there is some similarity in the way feed and permeate
are handled. The more common cassette has long been popular in the
pharmaceutical and biotechnical field. It too is a stack of flat-sheet
membranes, but the membrane is usually connected so that the feed
flows across the membrane elements in series to achieve higher con-
version per pass. Their popularity stems from easy direct scale-up from
laboratory to plant-scale equipment. Their limitation is that fluid man-
agement is inherently very limited and inefficient. Both types of cas-
sette are very compact and capable of automated manufacture.
Representative Process Applications
Pharmaceutical Removal of suspended matter is a frequent
application for MF. Processes may be either clarification, in which the
main product is a clarified liquid, or solids recovery. Separating cells
or their fragments from broth is the most common application. Clari-
fication of the broth in preparation for product recovery is the usual
objective, but the primary goal may be recovery of cells. Cross-flow
microfiltration competes well with centrifugation, conventional filtra-
tion by rotary vacuum filter or filter press and decantation. MF deliv-
ers a cleaner permeate, an uncontaminated, concentrated cell product
20-56 ALTERNATIVE SEPARATION PROCESSES
which may be washed in the process, and generally gives high yields.
There is no filter-aid disposal problem. Microfiltration has higher cap-
ital costs than the other processes, although total cost may be lower.
The recovery of penicillin is an example of a process for which cross-
flow microfiltration is generally accepted.
Water and Wastewater Microfiltration is beginning to be
applied to large-scale potable-water treatment. Its major advantage is
positive removal of cryptosporidium and giardia cysts, and its major
disadvantage is cost. MF is used in a few large sewage-treatment facil-
ities, where its primary advantage is that it permits a major reduction
in the physical size of the facility.
Chemical MF is used in several applications to recover caustic
values from cleaning or processing streams. An example is the caustic
solution used to clean dairy evaporators, which may be cleaned for
reuse by passing it through a microfilter. Significant savings in caustic
purchase and disposal costs provide the incentive. Acids are also
recovered and reused. Ceramic microfilters are most commonly used
in these applications.
Food and Dairy Microfiltration has many applications in the
food and dairy industries. An innovative dairy application uses MF
membranes to remove bacteria as a nonthermal means of disinfection
for milk. A special flow apparatus maintains a carefully controlled
transmembrane pressure as the milk flows across the membrane. The
concentrate contains the bacteria and spores, as well as any fat. The
concentrate may be heat sterilized and recombined with the sterile
permeate. In another milk application, some success is reported in
separating fat from milk or other dairy streams by cross-flow microfil-
tration instead of centrifugation. Transmembrane pressure must be
kept very low to prevent fat penetration into the membrane. In the
food industry, MF membranes are replacing diatomaceous earth fil-
tration in the processing of gelatin. The gelatin is passed with the per-
meate, but the haze producing components are retained. UF may be
used downstream to concentrate the gelatin.
Flow Schemes The outline of batch, semibatch, and stages-in-
series is given in the section describing ultrafiltration. Diafiltration is
also described there. All these techniques are common in MF, except
for stages-in-series, used rarely. MF features uses of special tech-
niques to control transmembrane pressure in some applications. An
example is one vendor’s device for the microfiltration of milk. In most
devices the permeate simply leaves by the nearest exit, but for this
application the permeate is pumped through the device in such a way
as to duplicate the pressure drop in the concentrate side, thus main-
taining a constant transmembrane-pressure driving force. In spite of
the low-pressure driving force, the flux is extremely high.
Limitations Some applications which seem ideal for MF, for
example the clarification of apple juice, are done with UF instead. The
reason is the presence of deformable solids which easily plug and
blind an MF membrane. The pores of an ultrafiltration membrane are
so small that this plugging does not occur, and high fluxes are main-
tained. UF can be used because there is no soluble macromolecule in
the juice that is desired in the filtrate. There are a few other significant
applications where MF seems obvious, but is not used because of
deformable particle plugging.
Economics Microfiltration may be the triumph of the Lillipu-
tians; nonetheless, there are a few large-industrial applications. Dex-
trose plants are very large, and as membrane filtration displaces the
precoat filters now standard in the industry, very large membrane
microfiltration equipment will be built.
Site Size Most MF processes require a smaller footprint than
competing processes. Reduction in total-area requirements are some-
times a decisive economic advantage for MF. It may be apparent that
the floor-space costs in a pharmaceutical facility are high, but munici-
pal facilities for water and sewage treatment are often located on
expensive real estate, giving MF an opportunity despite its higher
Large Plants The economics of microfiltration units costing
is treated under ultrafiltration. When ceramic membranes
are used, the cost optimum may shift energy consumption upward to
as much as 10 kWh/m3
Disposables For smaller MF applications, short membrane life is
a traditional characteristic. In these applications, costs are dominated
by the disposables, and an important characteristic of equipment
design is the ease, economy, and safety of membrane replacement.
Hygiene and Regulation Almost unique to MF is the influence
of regulatory concerns in selection and implementation of a suitable
microfilter. Since MF is heavily involved with industries regulated by
the Food and Drug Administration, concerns about process stability,
consistency of manufacture, virus reduction, pathogen control, and
material safety loom far larger than is usually found in other mem-
Process Description Gas-separation membranes separate gases
from other gases. Some gas filters, which remove liquids or solids
from gases, are microfiltration membranes. Gas membranes generally
work because individual gases differ in their solubility and diffusivity
through nonporous polymers. A few membranes operate by sieving,
Knudsen flow, or chemical complexation.
Selective gas permeation has been known for generations, and the
early use of palladium silver-alloy membranes achieved sporadic
industrial use. Gas separation on a massive scale was used to separate
using porous (Knudsen flow) membranes. An upgrade
of the membranes at Oak Ridge cost $1.5 billion. Polymeric mem-
branes became economically viable about 1980, introducing the mod-
ern era of gas-separation membranes. H2 recovery was the first major
application, followed quickly by acid gas separation (CO2 /CH4) and
the production of N2 from air.
The more permeable component is called the fast gas, so it is the
one enriched in the permeate stream. Permeability through polymers
is the product of solubility and diffusivity. The diffusivity of a gas in a
membrane is inversely proportional to its kinetic diameter, a value
determined from zeolite cage exclusion data (see Table 20-26 after
Breck, Zeolite Molecular Sieves, Wiley, New York, 1974, p. 636).
Tables 20-27, 20-28, and 20-29 provide units conversion factors useful
for calculations related to gas-separation membrane systems.
Leading Examples These applications are commercial, some on
a very large scale. They illustrate the range of application for gas-
separation membranes. Unless otherwise specified, all use polymeric
Hydrogen Hydrogen recovery was the first large commercial
membrane gas separation. Polysulfone fiber membranes became
available in 1980 at a time when H2 needs were rising, and these novel
membranes quickly came to dominate the market. Applications
include recovery of H2 from ammonia purge gas, and extraction of H2
from petroleum cracking streams. Hydrogen once diverted to low-
quality fuel use is now recovered to become ammonia, or is used to
desulfurize fuel, etc. H2 is the fast gas.
Carbon Dioxide-Methane Much of the natural gas produced in
the world is coproduced with an acid gas, most commonly CO2 and/or
H2S. While there are many successful processes for separating the
gases, membrane separation is a commercially successful competitor,
especially for small installations. The economics work best for feeds
with very high or very low CH4 content. Methane is a slow gas; CO2,
H2S, and H2O are fast gases.
Oxygen-Nitrogen Because of higher solubility, in many poly-
mers, O2 is faster than N2 by a factor of 5. Water is much faster still.
Since simple industrial single-stage air compressors provide suffi-
cient pressure to drive an air-separating membrane, moderate purity
N2 (95–99.5%) may be produced in low to moderate quantities quite
MEMBRANE SEPARATION PROCESSES 20-57
TABLE 20-26 Kinetic Diameters for Important Gases
Penetrant He H2 NO CO2 Ar O2 N2 CO CH4 C2H4 Xe C3H8
Kinetic dia, nm 0.26 0.289 0.317 0.33 0.34 0.346 0.364 0.376 0.38 0.39 0.396 0.43
economically by membrane separation. (Argon is counted as part of
the nitrogen.) An O2-enriched stream is a coproduct, but it is rarely of
economic value. The membrane process to produce O2 as a primary
product has a limited market.
Helium Helium is a very fast gas, and may be recovered from nat-
ural gas through the use of membranes. More commonly, membranes
are used to recover He after it has been used and become diluted.
Gas Dehydration Water is extremely permeable in polymer
membranes. Dehydration of air and other gases is a growing mem-
Vapor Recovery Organic vapors are recovered from gas streams
using highly permeable rubbery polymer membranes which are gen-
erally unsuitable for permanent gas separations because of poor selec-
tivity. The high sorption of vapors in these materials makes them ideal
for stripping and recovering vapors from gases.
Competing Processes Membranes are not the only way to make
these separations, neither are they generally the dominant way. In
many applications, membranes compete with cryogenic distillation
and with pressure-swing adsorption; in others, physical absorption is
the dominant method. The growth rate for membrane capacity is
higher than that for any competitor.
Basic Principles of Operation Gas-separation literature often
uses nomenclature derived from distillation, a practice that will
generally be followed here. L is the molar feed rate, V is the molar
permeate rate, R is the molar residue (L Ϫ V). Mole fractions of com-
ponents i, j, in the feed-residue phase will be xi, xj . . . and in the per-
meate phase yi, yj . . . . Stage cut, Θ, is permeate volume/feed volume,
Basic Equations In “Background and Definitions,” the basic
equation for gas permeation was derived with the major assumptions
Ji ∼ Ni = (ρi /z)(pi,feed − pi,permeate) (20-85)
where ρi is the permeability of component i through the membrane, Ji
is the flux of component i through the membrane for the partial pres-
sure difference (∆p) of component i. z is the effective thickness of the
membrane. By choosing units appropriately, J = ρ∆p.
A similar equation may be written for a second component, j, and
any additional number of components, employing partial pressures:
Jj = (ρj /z)(pj,feed − pj,permeate) (20-86)
The total pressure is the sum of the partial pressures:
Pfeed = Α (pfeed)i, j,... (20-87)
Ppermeate = Α (ppermeate)i, j... (20-88)
For simplicity, consider a two-component system. The volume frac-
tion of a component is
xi = (20-89)
yi = (20-90)
When two species are permeating through a membrane, the ratio of
their fluxes can be written following Eqs. (20-85) and (20-86) as:
Recalling Eq. (20-55), and restating it in the nomenclature for gas
α = (20-92)
Defining the pressure ratio Φ = Pfeed /Ppermeate and applying Eqs. (20-
87) through (20-90) gives:
= α΄ ΅ (20-93)
Combining these equations and rearranging, the permeate composi-
tion may be solved explicitly:
yi = ΄xi + + − Ίxi ++
Equation (20-94) gives a permeate concentration as a function of the
feed concentration at a stage cut, Θ = 0, To calculate permeate com-
position as a function of Θ, the equation may be used iteratively if the
permeate is unmixed, such as would apply in a test cell. The calcula-
tion for real devices must take into account the fact that the driving
force is variable due to changes on both sides of the membrane, as
partial pressure is a point function, nowhere constant. Using the same
caveat, permeation rates may be calculated component by component
using Eq. (20-86) and permeance values. For any real device, both
concentration and permeation require iterative calculations depen-
dent on module geometry.
Driving Force Gas moves across a membrane in response to a
difference in chemical potential. Partial pressure is sufficiently pro-
portional to be used as the variable for design calculations for most
gases of interest, but fugacity must be used for CO2 and usually for H2
at high pressure. Gas composition changes as a gas passes along a
membrane. As the fast gas passes through the membrane, xi drops.
Total pressure on the upstream side of the membrane drops because
of frictional losses in the device. Frictional losses on the permeate side
will affect the permeate pressure. The partial pressure of a compo-
nent in the permeate may thus rise rapidly. Permeation rate is a point
function, dependent on the difference in partial pressures at a point
on the membrane. Many variables affect point partial pressures,
among them are membrane structure, module design, and permeate
gas-sweep rates. Juxtaposition of feed and permeate is a function of
permeator design, and a rapid decline in driving force may result
when it is not expected (see “Membrane System Design Features”).
(α − 1)Φ
α − 1
α − 1
xi − (yi /Φ)
xj − (yj /Φ)
ᎏ (pi,feed − pi,permeate)
ᎏ (pj,feed − pj,permeate)
20-58 ALTERNATIVE SEPARATION PROCESSES
TABLE 20-27 Gas-Permeation Units
Quantity Engineering units Literature units SI units
Permeation rate Standard cubic feet/minute kmol/s
Permeation flux ft3
⋅day cm/sec (STP) kmol/m2
⋅day⋅psi Barrers kmol/m⋅s⋅Pa
⋅day⋅psi Barrers/cm kmol/m2
TABLE 20-28 Barrer Conversion Factors
Quantity Multiply By To get
Permeability Barrers 3.348 × 10−19
Permeability Barrers 4.810 × 10−8
Permeance Barrers/cm 3.348 × 10−17
Permeance Barrers/cm 1.466 × 10−6
TABLE 20-29 Industry-Specific Gas Measures
Cubic feet per
Industry-unit How measured pound mole kmol per mscf
STP, Mscf 1000 ft3
at 32°F 359.3 1.262
Gas industry, Mscf 1000 ft3
at 60°F 379.8 1.194
Air industry, Mnsf 1000 ft3
at 70°F 387.1 1.172
An additional complication may arise in a few cases from the Joule-
Thompson effect during expansion of a gas through a membrane
changing the temperature. High-pressure CO2 is an example.
Plasticization Gas solubility in the membrane is one of the fac-
tors governing its permeation, but the other factor, diffusivity, is not
always independent of solubility. If the solubility of a gas in a polymer
is too high, plasticization and swelling result, and the critical structure
that controls diffusion selectivity is disrupted. These effects are par-
ticularly troublesome with condensable gases, and are most often
noticed when the partial pressure of CO2 or H2S is high. H2 and He do
not show this effect This problem is well known, but its manifestation
is not always immediate.
Limiting Cases Equation (20-94) has two limiting cases for a
binary system. First, when α ӷ φ. In this case, selectivity is no longer
yi Х xi = xiΦ (20-95)
Module design is very important for this case, as the high α may result
in high permeate partial pressure. An example is the separation of
H2O from air.
Conversely, when Φ ӷ α, pressure ratio loses its importance, and
permeate composition is:
yi Х (20-96)
Module design for this case is of lesser importance.
Selectivity and Permeability
State of the Art A desirable gas membrane has high separating
power (α) and high permeability to the fast gas, in addition to critical
requirements discussed below. The search for an ideal membrane
produced copious data on many polymers, neatly summarized by
Robeson [J. Membrane Sci., 62, 165 (1991)]. Plotting log permeabil-
ity versus log selectivity (α), an “upper bound” is found (see Fig.
20-69) which all the many hundreds of data points fit. The data were
taken between 20 and 50°C, generally at 25 or 35°C.
The lower line in Fig. 20-69 shows the upper bound in 1980.
Although no breakthrough polymers have been reported in the past
[1 + xi(α − 1)]
few years, it would be surprising if these lines remain the state of the
art forever. Performance parameters for polymers used for O2/N2 sep-
arations are given in Table 20-30.
The upper-bound line connects discontinuous points, but polymers
exist near the bound for separations of interest. Whether these will be
available as membranes is a different matter. A useful membrane
requires a polymer which can be fabricated into a device having an active
layer around 50 nm thick. At this thickness, membrane properties may
vary significantly from bulk properties, although not by a factor of 2.
The data reported are permeabilities, not fluxes. Flux is propor-
tional to permeability/thickness. The separations designer must deal
with real membranes, for which thickness is determined by factors
outside the designer’s control. It is vital that flux data are used in
Glassy polymers are significantly overrepresented in the high-
selectivity region near the upper bound, and rubbery polymers are
overrepresented at the high-permeability end, although the highest-
permeability polymer discovered, poly(trimethylsilylpropyne) is glassy.
Most of the polymers with interesting properties are noncrystalline.
Current membrane-materials research is strongly focused on glassy
materials and on attempts to improve diffusivity, as it seems more
promising than attempts to increase solubility [Koros, North Am.
Membrane Soc. Sixth Annual Meeting, Breckenridge, CO (1994)].
Robeson [J. Membrane Sci., 62, 165 (1991); Polymer, 35, 4970
(1994)] has determined upper-bound lines for many permeant pairs
in hundreds of polymers. These lines may be drawn from Eq. (20-97)
and the data included in Table 20-31. These values will give ρi in
MEMBRANE SEPARATION PROCESSES 20-59
FIG. 20-69 Plot of separation factor versus permeability for many polymers, O2/N2. Abscissa—“Fast Gas Permeability, ρ(O2)
Barrers.” Ordinate—“Selectivity, α (O2/N2).”
TABLE 20-30 High-Performance Polymers for O2/N2*
Polymer α (O2/N2) ρ (O2)—Barrers
Poly (trimethylsilylpropyne) 2.0 4000
Tetrabromo Bis A polycarbonate 7.47 1.36
Poly (tert-butyl acetylene) 3.0 300
Vectra polyester 15.3 0.00046
Poly (triazole) 9.0 1.2
Polypyrrolone 6.5 7.9
*Polymers that are near the upper bound and their characteristics.
Barrers; α is dimensionless. Robeson [op.cit., (1991); op. cit., (1994)]
lists high-performance polymers for most of these gas pairs, like Table
log ρi = log k − m log αij (20-97)
Temperature Effects A temperature increase in a polymer
membrane permits larger segmental motions in the polymer, produc-
ing a dramatic increase in diffusivity. Countering this is a decrease in
solubility. It increases the size of the gaps in the polymer matrix,
decreasing diffusivity selectivity. The net result is that for a glassy
polymer, permeability rises while selectivity declines. For organic per-
meants in rubbery polymers, this trend is often reversed.
Plasticization and Other Time Effects Most data from the lit-
erature, including those presented above are taken from experiments
where one gas at a time is tested, with α calculated as a ratio of the two
permeabilities. If either gas permeates because of a high-sorption
coefficient rather than a high diffusivity, there may be an increase in
the permeability of all gases in contact with the membrane. Thus, the
α actually found in a real separation may be much lower than that cal-
culated by the simple ratio of permeabilities. The data in the literature
do not reliably include the plasticization effect. If present, it results in
the sometimes slow relaxation of polymer structure giving a rise in
permeability and a dramatic decline in selectivity.
Other Caveats Transport in glassy polymers is different from
transport in rubbery ones. In glassy polymers, there are two sites in
which sorption may occur, and the literature dealing with dual-mode
sorption is voluminous. The simplest case describes behavior when
the downstream pressure is zero. It is of great help in understanding
the theory but of limited value in practice. There are concerns about
permeation of mixtures in glassy polymers with reports of crowding
out and competitive sorption. Practical devices are built and operated
for many streams, and the complications are often minor. But taking
data independently determined for two pure gases and dividing them
to obtain α in the absence of other facts is risky.
Organic Organic polymer membranes are the basis for almost all
commercial gas-separation activity. Early membranes were cellulose
esters and polysulfone. These membranes have a large installed base.
New installations are dominated by specialty polymers designed for
the purpose, including some polyimides and halogenated polycarbon-
ates. In addition to skinned membranes, composites are made from
“designer” polymers, requiring as little as 2 g/1000 m2
. The rapid rise
of N2 /O2 membranes in particular is the result of stunning improve-
ments in product uniformity and quality. A few broken fibers in a
module results in the module’s being scrapped.
Caulked Membrane manufacturing defects are unavoidable, and
pinholes are particularly deleterious in gas-separation membranes. A
very effective remedy is to caulk the membrane by applying a highly
permeable, very thin topcoat over the finished membrane. While the
coating will have poor selectivity, it will plug up the gross leaks while
impairing the fast gas permeance only slightly. Unless the as-cast
membrane is almost perfect, caulking dramatically improves mem-
Metallic Palladium films pass H2 readily, especially above 300°C.
α for this separation is extremely high, and H2 produced by purifica-
tion through certain Pd alloy membranes is uniquely pure. Pd alloys
are used to overcome the crystalline instability of pure Pd during heat-
ing-cooling cycles. Economics limit this membrane to high-purity
Advanced Materials Experimental membranes have shown re-
markable separations between gas pairs such as O2 /N2 whose kinetic
diameters (see Table 20-26) are quite close. Most prominent is the
carbon molecular sieve membrane, which operates by ultramicro-
porous molecular sieving. Preparation of large-scale permeators based
on ultramicroporous membranes has proven to be a major challenge.
Catalytic A catalytic-membrane reactor is a combination hetero-
geneous catalyst and permselective membrane that promotes a reac-
tion, allowing one component to permeate. Many of the reactions
studied involve H2. Membranes are metal (Pd, Ag), nonporous metal
oxides, and porous structures of ceramic and glass. Falconer, Noble,
and Sperry [in Noble and Stern (eds.), op. cit., pp. 669–709] review
status and potential developments.
Membrane System Design Features For the rate process of
permeation to occur, there must be a driving force. For gas separa-
tions, that force is partial pressure (or fugacity). Since the ratio of the
component fluxes determines the separation, the partial pressure of
each component at each point is important. There are three ways of
driving the process: Either high partial pressure on the feed side
(achieved by high total pressure), or low partial pressure on the per-
meate side, which may be achieved either by vacuum or by introduc-
tion of a sweep gas. Both of the permeate options have negative
economic implications, and they are less commonly used.
Figure 20-70 shows three of the principal operating modes for gas
membranes. A critical issue is the actual partial pressure of permeant
at a point on the membrane. Flow arrangements for the permeate are
very important in determining the efficiency of the separation, in
rough analogy to the importance of arrangements in heat exchangers.
Spiral membranes are the usual way to form flat sheet into modules.
They have the characteristic that the feed and the permeate move at
20-60 ALTERNATIVE SEPARATION PROCESSES
TABLE 20-31 Upper-Bound
Coordinates for Gas Pairs
Gas pair log k m
He/N2 4.0969 1.0242
H2 /N2 4.7236 1.5275
He/CH4 3.6991 0.7857
H2 /CH4 4.2672 1.2112
O2 /N2 5.5902 5.800
He/O2 3.6628 1.295
H2 /O2 4.5534 2.277
CO2 /CH4 6.0309 2.6264
He/H2 2.9823 4.9535
He/CO2 2.8482 1.220
H2 /CO2 3.0792 1.9363
FIG. 20-70 Flow paths in gas permeators. (Courtesy Elsevier.)
right angles. Since the membrane is always cast on a porous support,
point-permeate values are influenced by the substrate.
Hollow-fiber membranes may be run with shell-side or tube-side
feed, cocurrent, countercurrent or in the case of shell-side feed and
two end permeate collection, co- and countercurrent. Not shown is
the scheme for feed inside the fiber, common practice in lower-
pressure separations such as air.
The design of the membrane device will influence whether the
membrane is operating near its theoretical limit. Sengupta and Sirkar
[in Noble and Stern (eds.), op. cit., pp. 499–552] treat module design
thoroughly (including numerical examples for most module configu-
rations) and provide an extensive bibliography.
For a hollow-fiber device running with shell-side feed with the
membrane on the outside, Giglia et al. [Ind. Eng. Chem Research, 29,
1239–1248 (1991)] analyzed the effect of membrane-backing porosity
on separation efficiency. The application is production of N2 from air,
the desired result being the lowest possible O2 content in the reten-
tate at a given stage cut. The modules used were operated cocurrent
and countercurrent. If the porous-membrane backing prevented the
permeate from mixing with the gas adjacent to the membrane, a result
approximating cross-flow is expected. For the particular membrane
structure used, the experimental result for cocurrent flow was quite
close to the calculated cocurrent value, while for countercurrent flow,
the experimental data were between the values calculated for the
crosscurrent model and the countercurrent model. For a membrane
to be commercially useful in this application, the mass transfer on the
permeate side must exceed the crosscurrent model.
Air is commonly run with tube-side feed. The permeate is run
countercurrent with the separating skin in contact with the permeate.
(The feed gas is in contact with the macroporous back side of the
membrane.) This configuration has proven to be superior, since the
permeate-side mass-transfer problem is reduced to a minimum, and
the feed-side mass-transfer problem is not limiting.
Partial Pressure Pinch An example of the limitations of the par-
tial pressure pinch is the dehumidification of air by membrane. While
O2 is the fast gas in air separation, in this application H2O is faster still.
Special dehydration membranes exhibit α = 20,000. As gas passes
down the membrane, the partial pressure of H2O drops rapidly in the
feed. Since the H2O in the permeate is diluted only by the O2 and N2
permeating simultaneously, pH2O rises rapidly in the permeate. Soon
there is no driving force. The commercial solution is to take some of
the dry air product and introduce it into the permeate side as a coun-
tercurrent sweep gas, to dilute the permeate and lower the H2O par-
tial pressure. It is in effect the introduction of a leak into the
membrane, but it is a controlled leak and it is introduced at the opti-
Fouling Industrial streams may contain condensable or reactive
components which may coat, solvate, fill the free volume, or react with
the membrane. Gases compressed by an oil-lubricated compressor
may contain oil, or may be at the water dew point. Materials that will
coat or harm the membrane must be removed before the gas is
treated. Most membranes require removal of compressor oil. The
extremely permeable poly(trimethylsilylpropyne) may not become a
practical membrane because it loses its permeability rapidly. Part of
the problem is pore collapse, but it seems extremely sensitive to con-
tamination even by diffusion pump oil and gaskets [Robeson, op. cit.,
A foulinglike problem may occur when condensable vapors are left
in the residue. Condensation may result which in the best case results
in blinding of the membrane, and in the usual case, destruction of the
membrane module. Dew-point considerations must be part of any
gas-membrane design exercise.
Modules and Housings Modern gas membranes are packaged
either as hollow-fiber bundles or as spiral-wound modules. The for-
mer uses extruded hollow fibers. Tube-side feed is preferable, but it is
limited to about 1.5 MPa. Higher-pressure applications are usually fed
on the shell side. A large industrial permeator contains fibers 400 µm
by 200 µm i.d. in a 6-inch shell 10 feet long. Flat-sheet membrane is
wound into spirals, with an 8- by 36-inch permeator containing 25 m2
of membrane. Both types of module are similar to those illustrated in
“Background and Definitions.” Spiral modules are useful when feed
flows are very high and especially in vapor-permeation applications.
Otherwise, fiber modules have a large and growing share of the
Energy Requirements The thermodynamic minimum energy
requirement to separate a metric ton of N2 from air and compress it to
atmospheric pressure at 25°C is independent of separation method,
20.8 MJ or 5.8 kWh. In practice, a cryogenic distillation plant requires
twice this energy, and it produces a very pure product as a matter of
course. The membrane process requires somewhat more energy than
distillation at low purity and much more energy at high purity. Mem-
branes for O2-reduced air are economical and are a rapidly growing
application, but not because of energy efficiency. Figure 20-71 may be
used to estimate membrane energy requirements. From the required
purity, locate the stage cut on either of the countercurrent flow
curves. The compressor work required is calculated using the pres-
sure and the term (product flow rate)/(1 − Θ). The N2-rich product is
produced at pressure, and the O2-enriched permeate is vented. For
example, if gas is required containing 97 percent inerts, assume the
product composition would be 96 percent N2, 2 percent O2, 1 percent
Ar, and 3 percent O2, giving a calculated molar mass of 28.2. One
tonne would thus contain 35.4 kmol. For 98 percent inerts, Fig. 20-71
shows a stage cut of 67 percent when operating at 690 kPa (abs) and
25°C. Therefore, 35.5/(1 − 0.67) = 108 kmol of air are required as
feed. The adiabatic compression of 108 kmol of air from atmospheric
pressure and 300 K (it would subsequently be cooled to 273 + 25 =
298 K) requires 192 kWh. Assuming 75 percent overall compressor +
driver efficiency, 256 kWh are required. For comparison, a very effi-
cient, large N2 distillation plant would produce 99.99 percent N2 at
690 kPa for 113 kWh/metric ton. The thermodynamic minimum to
separate (N2 + Ar) and deliver it at the given P and T is 72 kWh.
Economics It is ironic that a great virtue of membranes, their
versatility, makes economic optimization of a membrane process very
difficult. Designs can be tailored to very specific applications, but
each design requires a sophisticated computer program to optimize its
costs. Spillman [in Noble and Stern (eds)., op. cit., pp. 589–667] pro-
vides an overall review and numerous specific examples including
circa 1989 economics.
Rules of Thumb With a few notable exceptions such as H2
through Pd membranes, membrane separations are not favored when
a component is required at high purity. Often, membranes serve these
needs by providing a moderate purity product which may be inexpen-
sively upgraded by a subsequent process. Increasing the purity of N2
by the introduction of H2 or CH4 to react with unwanted O2 is a good
example. Unless permeates are recycled, high product purity is
accompanied by lower product recovery.
Pressure ratio (Φ) is quite important, but transmembrane ∆P mat-
ters as well. Consider the case of a vacuum permeate (Φ = ∞): The
membrane area will be an inverse function of P. Φ influences separa-
tion and area, transmembrane ∆P influences area.
In a binary separation, the highest purity of integrated permeate
occurs at Θ = 0. Purity decreases monotonically until it reaches the
feed purity at Θ = 1. In a ternary system, the residue concentration of
the gas with the intermediate permeability will reach a maximum at
some intermediate stage cut.
Concentration polarization is a significant problem only in vapor sep-
aration. There, because the partial pressure of the penetrant is normally
low and its solubility in the membrane is high, there can be depletion in
the gas phase at the membrane. In other applications it is usually safe to
assume bulk gas concentration right up to the membrane.
Another factor to remember is that for α = 1, or for Φ = 1, or for
(1/Θ) = 1 there is no separation at all. Increasing any of the quantities
as defined make for a better separation, but the improvement is
diminishing in all cases as the value moves higher. An example of the
economic tradeoff between permeability and α is illustrated in Fig.
20-72 where the economics are clearly improved by sacrificing selec-
tivity for flux.
Compression If compression of either feed or permeate is
required, it is highly likely that compression capital and operating
costs will dominate the economics of the gas-separation process.
In some applications, pressure is essentially “free,” such as when
removing small quantities of CO2 from natural gas. The gas is often
MEMBRANE SEPARATION PROCESSES 20-61
produced at pressure, but is compressed for transmission anyway, and
since the residue constitutes the product, it continues downstream at
pressure. H2 frequently enters the separation process at pressure, an
advantage for membranes. Unlike CH4, the H2 is in the permeate, and
recompression may be a significant cost. A relative area cure for
CO2 /CH4 is shown in Fig. 20-73. When the permeate is the product
(H2, CO2) the increasing membrane area shown in Fig. 20-74 is largely
the effect of more gas to pass through the membrane, since the curve
is based on a constant volume of feed gas, not a constant output of H2.
The facts of life in compressor economics are in painful opposition to
the desires of the membrane designer. Pressure ratios higher than six
become expensive; vacuum is very expensive, and scale is important.
Because of compressor economics, staging membranes with recom-
pression is unusual. Designers can assume that a flow sheet that mixes
unlike streams or reduces pressure through a throttling valve will
increase cost in most cases.
Product Losses Account must be taken of the product loss, the
slow gas in the permeate (such as CH4), or the fast gas in the residue
(such as H2). Figure 20-75 illustrates the issue for a membrane used to
purify natural gas from 93 percent to pipeline quality, 98 percent. In
the upper figure, the gas is run through a permeator bank operating as
a single stage. For the membrane and module chosen, the permeate
contains 63 percent CH4. By dividing the same membrane area into
two stages, two permeates (or more) may be produced, one of which
20-62 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-71 Air fractionation by membrane. O2 in retentate as a function of feed fraction passed through the mem-
brane (stage cut) showing the different result with changing process paths. Process has shell-side feed at 690 kPa
(abs) and 298 K. Module comprised of hollow fibers, diameter 370 µm od × 145 µm id × 1500 mm long. Membrane
properties α = 5.7 (O2 /N2), permeance for O2 = 3.75 × 10−6
Barrer/cm. (Courtesy Innovative Membrane Systems/
may have significantly higher economic value than the single mixed
permeate. In fact, where CH4 is involved, another design parameter is
the local economic value of various waste streams as fuel.
Membrane Replacement Membrane replacement is a signifi-
cant cost factor, but membrane life and reliability are now reasonable.
Membranes are more susceptible to operating upsets than more tra-
ditional equipment, but their field-reliability record in properly engi-
neered, properly maintained installations is good to excellent. In N2
separations, membrane life is very long.
Competing Technologies The determination of which separa-
tion technique is best for a specific application is a dynamic function
of advances in membranes and several other technologies. At this
writing, very small quantities of pure components are best obtained by
purchase of gas in cylinders. For N2, membranes become competitive
at moderate flow rates where purity required is <99.5 percent. At
higher flow rates and higher purity, pressure swing adsorption (PSA)
is better (see Sec. 16: “Adsorption and Ion Exchange”). At still higher
volumes, delivered liquid N2, pipeline N2, or on-site distillation will be
superior. For O2, membranes have little economic importance. The
problem is the economic cost of using vacuum on the permeate side,
or the equally unattractive prospect of compressing the feed and oper-
ating at low stage cut. For H2, membranes are dominant when the
feed is at high pressure, except for high purity (excepting Pd noted
above) and very large volume. Higher purity or lower feed pressure
favor PSA. Very high volume favors cryogenic separation. For CH4,
membranes compete at low purity and near pipeline (98 percent)
purity, but distillation and absorption are very competitive at large
scale and for intermediate CO2 contamination levels.
Membranes are found as adjuncts to most conventional processes,
since their use can improve overall economics in cases where mem-
brane strength coincides with conventional process weakness.
Process Description Pervaporation is a separation process in
which a liquid mixture contacts a nonporous permselective mem-
brane. One component is transported through the membrane prefer-
entially. It evaporates on the downstream side of the membrane
leaving as a vapor. The name is a contraction of permeation and evap-
oration. Permeation is induced by lowering partial pressure of the per-
meating component, usually by vacuum or occasionally with a sweep
gas. The permeate is then condensed or recovered. Thus, three steps
are necessary: Sorption of the permeating components into the mem-
brane, diffusive transport across the nonporous membrane, then de-
sorption into the permeate space, with a heat effect. Pervaporation
membranes are chosen for high selectivity, and the permeate is often
MEMBRANE SEPARATION PROCESSES 20-63
FIG. 20-72 Constant-cost lines as a function of permeability and selectivity for
CO2 /CH4. Cellulose-acetate membrane “mscf” is one thousand standard cubic
feet. (Courtesy W. R. Grace.)
FIG. 20-73 Influence of feed purity on total membrane area when the residue
gas at fixed purity is the product. Feed-gas volume is constant. CO2 /CH4 cellu-
lose-acetate membrane, α = 21. (Courtesy W. R. Grace.)
FIG. 20-74 Influence of feed purity on total membrane area when the per-
meate gas at fixed purity is the product. Feed-gas volume is constant. H2 /CH4
cellulose-acetate membrane, α = 45. (Courtesy W. R. Grace.)
FIG. 20-75 Effect on permeate of dividing a one-stage separation into two
equal stages having the same total membrane area. Compositions of A, D, and F
are equal for both cases. (Courtesy W. R. Grace.)
In the flow schematic (Fig. 20-76), the condenser controls the
vapor pressure of the permeating component. The vacuum pump, as
shown, pumps both liquid and vapor phases from the condenser. Its
major duty is the removal of noncondensibles. Early work in pervapo-
ration focused on organic-organic separations. Many have been
demonstrated; few if any have been commercialized. Still, there are
prospects for some difficult organic separations.
An important characteristic of pervaporation that distinguishes it
from distillation is that it is a rate process, not an equilibrium process.
The more permeable component may be the less-volatile component.
Pervaporation has its greatest utility in the resolution of azeotropes, as
an adjunct to distillation. Selecting a membrane permeable to the
minor component is important, since the membrane area required is
roughly proportional to the mass of permeate. Thus pervaporation
devices for the purification of the ethanol-water azeotrope (95 per-
cent ethanol) are always based on a hydrophilic membrane.
Pervaporation membranes are of two general types. Hydrophilic
membranes are used to remove water from organic solutions, often
from azeotropes. Hydrophobic membranes are used to remove
organic compounds from water. The important operating characteris-
tics of hydrophobic and hydrophilic membranes differ. Hydrophobic
membranes are usually used where the solvent concentration is about
6 percent or less—above this value, other separation methods are usu-
ally cheaper unless the flow rate is small. At low-solvent levels the
usual membrane (silicone rubber) is not swollen appreciably, and
movement of solvent into the membrane makes depletion of solvent
in the boundary layer (concentration polarization) an important
design problem. Hydrophilic pervaporation membranes operate such
that the upstream portion is usually swollen with water, while the
downstream face is low in water concentration because it is being
depleted by vaporization. Fluxes are low enough (<5 kg/m2
boundary layer depletion (liquid side) is not limiting.
The simplifying assumptions that make Fick’s law useful for other
processes are not valid for pervaporation. The activity gradient across
the membrane is far more important than the pressure gradient.
Equation (20-98) is generally used to describe the pervaporation
Ji = −DiC (20-98)
where ai is the activity coefficient of component i.
This equation is not particularly useful in practice, since it is diffi-
cult to quantify the relationship between concentration and activity.
The Flory-Huggins theory does not work well with the crosslinked
semicrystalline polymers that comprise an important class of per-
vaporation membranes. Neel (in Noble and Stern, op. cit., pp. 169–
176) reviews modifications of the Stefan-Maxwell approach and other
equations of state appropriate for the process.
d ln ai
A typical permeant-concentration profile in a pervaporation mem-
brane is shown in Fig. 20-77. The concentration gradient at the per-
meate (vapor side) of the hydrophilic membrane is usually rate
controlling. Therefore the downstream pressure (usually controlled
by condenser temperature) is very important for flux and selectivity.
Since selectivity is the ratio of fluxes of the components in the feed, as
downstream pressure increases, membrane swelling at the permeate
interface increases, and the concentration gradient at the permeate
interface decreases. The permeate flux drops, and the more swollen
membrane is less selective. A rise in permeate pressure may result in
a drastic drop in membrane selectivity. This effect is diminished at low
water concentrations, where membrane swelling is no longer domi-
nant. In fact, when the water concentration drops far enough, perme-
ate backpressure looses its significance. (See Fig. 20-78.)
For rubbery membranes (hydrophobic), the degree of swelling has
less effect on selectivity. Thus the permeate pressure is less critical
to the separation, but it is critical to the driving force, thus flux, since the
vapor pressure of the organic will be high compared to that of water.
Definitions Following the practice presented under “Gas-
Separation Membranes,” distillation notation is used. Literature arti-
cles often use mass fraction instead of mole fraction, but the
substitution of one to the other is easily made.
β ϵ (20-99)
20-64 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-76 Simplified flow schematic for a pervaporation system. Heated feed enters from left through a feed
pump. Heaters in a recirculating feed loop may be required (not shown). Stripped liquid exits at the top of the per-
vaporation membrane. Vapor exits at the bottom to a condenser. Liquid and noncondensibles are removed under vac-
uum. (Courtesy Hoechst Celanese.)
FIG. 20-77 Permeant-concentration profile in a pervaporation membrane. 1—
Upstream side (swollen). 2—Convex curvature due to concentration-dependent
permeant diffusivity. 3—Downstream concentration gradient. 4—Exit surface of
membrane, depleted of permeant, thus unswollen. (Courtesy Elsevier.)
where β is the enrichment factor. β is related to α [Eq. (20-92)], the
separation factor, by
β = α = (20-100)
α is larger than β, and conveys more meaning when the membrane
approaches the ideal. β is preferred in pervaporation because it is eas-
ier to use in formulations for cost, yield, and capacity. In fact, note that
neither factor is constant, and that both generally change with xi and
Operational Factors In industrial use, pervaporation is a
continuous-flow single-stage process. Multistage cascade devices are
unusual. Pervaporation is usually an adjunct separation, occasionally a
principal one. It is used either to break an azeotrope or to concentrate
a minor component. Large stand-alone uses may develop in areas
where distillation is at a disadvantage, such as with closely boiling
components, but these developments, have not yet been made.
Notable exceptions occur when a stream is already fairly pure, such as
a contaminated water source, or isopropanol from microelectronics
fabrication washing containing perhaps 15 percent water.
In continuous flow, the feed will begin with concentration (xi,feed)
and be stripped until it reaches some desired (xi,residue). If significant
mass is removed from the feed, it will cool, since the general rule for
liquids is that the latent heat exceeds the specific heat per °C by two
orders of magnitude (more for water). So the membrane has a con-
centration gradient and a temperature gradient both across it and
along it. Pervaporation is a complex multigradient process.
The term β is not constant for some important separations. Even
worse, it can exhibit maxima that make analytic treatment difficult.
Operating diagrams are often used for preliminary design rather than
equations. Because of the very complex behavior in the membrane as
concentrations change, all design begins with an experiment. For water
removal applications, design equations often mispredict the rate con-
stants as the water content of the feed approaches zero. The estimates
tend to be lower than the experimental values, which would lead to
overdesign. Therefore it is necessary to obtain experimental data over
the entire range of water concentrations encountered in the separa-
tion. Once the kinetic data are available, heat transfer and heat capac-
ity are the problem. It is general practice to pilot the separation on a
prototype module to measure the changes due to thermal effects. This
is particularly true for water as the permeant, given its high latent heat.
For removal of an organic component from water, swelling of
the organophillic membrane would result in a higher flux and lower α.
At organic levels below about 10 percent, that has not been a major
β(1 − xi)
1 − xiβ
1 + (α − 1)xi
problem. In most applications, boundary-layer mass-transfer limitations
become limiting. Pilot data must therefore be taken with hydrodynamic
similarity to the module that will be used and the actual organic perme-
ation rate may become limited by the boundary layer. In organic-
organic pervaporation, membrane swelling is a major concern.
Vapor Feed A variant on pervaporation is to use vapor, rather
than liquid, as a feed. While the resulting process could be classified
along with gas-separation membranes, it is customarily regarded as
The residue at the top of a distillation column is a vapor, so there is
logic in using it as the feed to a membrane separator. Weighed against
the obvious advantage are disadvantages of vapor handling (compres-
sors cost more than pumps) and equipment size to handle the larger
vapor volume. When the more volatile component permeates, heat
must be added to maintain superheat. The vapor feed technique has
been used in a few large installations where the advantages outweigh
Dehydration The growing use of isopropanol as a clean-rinse
fluid in microelectronics produces significant quantities of an 85–90
percent isopropanol waste. Removing the water and trace contami-
nants is required before the alcohol can be reused. Pervaporation pro-
duces a 99.99 percent alcohol product in one step. It is subsequently
polished to remove metals and organics. In Europe, dehydration of
ethanol is the largest pervaporation application. For the very large
ethanol plants typical of the United States, pervaporation is not com-
petitive with thermally integrated distillation.
Organic from Water An area where pervaporation may become
important is in flavors, fragrances, and essential oils. Here, high-value
materials with unique properties are recovered from aqueous or alco-
Pollution Control Pervaporation is used to reduce the organic
loading of a waste stream, thus effecting product recovery and a reduc-
tion in waste-treatment costs. An illustration is a waste stream contain-
ing 11 percent (wt) n-propanol. The residue is stripped to 0.5 percent
and 96 percent of the alcohol is recovered in the permeate as a 45 per-
cent solution. This application uses a hydrophobic, rubbery mem-
brane. The residue is sent to a conventional waste-treatment plant.
Pervaporation Membranes Pervaporation has a long history,
and many materials have found use in pervaporation experiments.
Cellulosic-based materials have given way to polyvinyl alcohol and
blends of polyvinyl alcohol and acrylics in commercial water-removing
membranes. These membranes are typically solution cast (from
MEMBRANE SEPARATION PROCESSES 20-65
FIG. 20-78 Pervaporation schematic for ethanol-water. The illustration shows the complex behavior for a sim-
ple system at three pressures. Only the region above 90 percent is of commercial interest. (Courtesy Elsevier.)
water) on ultrafiltration membrane substrates. It is important to have
enough crosslinking in the final polymer to avoid dissolution of the
membrane in use. A very thin membrane with little penetration into
the UF substrate is required. The substrate can be a problem, as it
provides significant pressure drop to the vapor passing through it
which, as mentioned above, has serious ramifications for flux and sep-
aration efficiency. Many new membranes are under development.
Ion-exchange membranes, and polymers deposited by or polymerized
by plasma, are frequently mentioned in the literature.
Modules Every module design used in other membrane opera-
tions has been tried in pervaporation. One unique requirement is for
low hydraulic resistance on the permeate side, since permeate pres-
sure is very low (0.1–1 Pa). The rule for near-vacuum operation is the
bigger the channel, the better the transport. Another unique need is
for heat input. The heat of evaporation comes from the liquid, and
intermediate heating is usually necessary. Of course economy is
always a factor. Plate-and-frame construction was the first to be used
in large installations, and it continues to be quite important. Some
smaller plants use spiral-wound modules, and some membranes can
be made as capillary bundles. The capillary device with the feed on
the inside of the tube has many advantages in principle, such as good
vapor-side mass transfer and economical construction, but it is still
limited by the availability of membrane in capillary form.
GENERAL REFERENCES: This section is based on three publications by Heiner
Strathmann, to which the interested reader is referred for greater detail [chap.
6, pp. 213–281, in Noble and Stern (eds.), op. cit.; sec. V, pp. 217–262, in Ho and
Sirkar, op. cit.; chap. 8, pp. 8-1–8-53, Baker et al., op. cit.].
Process Description Electrodialysis (ED) is a membrane sepa-
ration process in which ionic species are separated from water,
macrosolutes, and all uncharged solutes. Ions are induced to move by
an electrical potential, and separation is facilitated by ion-exchange
membranes. Membranes are highly selective, passing either anions or
cations and very little else. The principle of ED is shown in Fig. 20-79.
The feed solution containing ions enters a compartment whose
walls are a cation-exchange and an anion-exchange membrane. If the
anion-exchange membrane is in the direction of the anode, as shown
for the middle feed compartment, anions may pass through that mem-
brane in response to an electrical potential. The cations can likewise
move toward the cathode. When the ions arrive in the adjacent com-
partments, however, their further progress toward the electrodes is
prevented by a membrane having the same electrical charge as the
ion. The two feed compartments to the left and right of the central
compartment are concentrate compartments. Ions entering these two
compartments, either in the feed or by passing through a membrane,
are retained, either by a same-charged membrane or by the EMF
driving the operation. The figure shows two cells (four membranes)
between anode and cathode. In an industrial application, a membrane
stack can be composed of hundreds of cells, where mobile ions are
alternately being depleted and concentrated.
Many related processes use charged membranes and/or EMF.
Electrodialytic water dissociation (water splitting), diffusion dialysis,
Donnan dialysis, and electrolysis are related processes. Electrolysis
(chlorine-caustic) is a process of enormous importance much of which
is processed through very special membranes.
Leading Examples Electrodialysis has its greatest use in remov-
ing salts from brackish water, where feed salinity is around 0.05–0.5
percent. For producing high-purity water, ED can economically
reduce solute levels to extremely low levels as a hybrid process in
combination with an ion-exchange bed. ED is not economical for the
production of potable water from seawater. Paradoxically, it is also
used for the concentration of seawater from 3.5 to 20 percent salt.
The concentration of monovalent ions and selective removal of diva-
lent ions from seawater uses special membranes. This process is
unique to Japan, where by law it is used to produce essentially all of
its domestic table salt. ED is very widely used for deashing whey,
where the desalted product is a useful food additive, especially for
Many ED-related processes are practiced on a small scale, or in
unique applications. Electrodialysis may be said to do these things
well: separate electrolytes from nonelectrolytes and concentrate elec-
trolytes to high levels. It can do this even when the pH is very low. ED
does not do well at: removing the last traces of salt (although the
hybrid process, electrodeionization, is an exception), running at high
pH, tolerating surfactants, or running under conditions where solubility
20-66 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-79 Schematic diagram of electrodialysis. Solution containing electrolyte is alternately depleted or
concentrated in response to the electrical field. Feed rates to the concentrate and diluate cells need not be
equal. In practice, there would be many cells between electrodes.
limits may be exceeded. Hydroxyl ion and especially hydrogen ion eas-
ily permeate both types of ED membrane. Thus, processes that gen-
erate a pH gradient across a membrane are limited in their scope.
Water splitting, a closely related process, is useful for reconstituting
an acid and a base out of a salt. It is used to reclaim salts produced
Membranes Ion-exchange membranes are highly swollen gels
containing polymers with a fixed ionic charge. In the interstices of the
polymer are mobile counterions. A schematic diagram of a cation-
exchange membrane is depicted in Fig. 20-80.
Figure 20-80 is a schematic diagram of a cation-exchange mem-
brane. The parallel, curved lines represent the polymer matrix com-
posed of an ionic, crosslinked polymer. Shown in the polymer matrix
are the fixed negative charges on the polymer, usually from sulfonate
groups. The spaces between the polymer matrix are the water-swollen
interstices. Positive ions are mobile in this phase, but negative ions are
repelled by the negative charge from the fixed charges on the polymer.
In addition to high permselectivity, the membrane must have low-
electrical resistance. That means it is conductive to counterions and
does not unduly restrict their passage. Physical and chemical stability
are also required. Membranes must be mechanically strong and
robust, they must not swell or shrink appreciably as ionic strength
changes, and they must not wrinkle or deform under thermal stress.
In the course of normal use, membranes may be expected to
encounter the gamut of pH, so they should be stable from 0 < pH < 14
and in the presence of oxidants.
Optimization of an ion-exchange membrane involves major trade-
offs. Mechanical properties improve with crosslink density, but so
does high electrical resistance. High concentration of fixed charges
favors low electrical resistance and high selectivity, but it leads to high
swelling, thus poor mechanical stability. Membrane developers try to
combine stable polymeric backbones with stable ionic functional
groups. The polymers are usually hydrophobic and insoluble. Poly-
styrene is the major polymer used, with polyethylene and polysulfone
finding limited application.
Most commercial ion-exchange membranes are homogeneous, pro-
duced either by polymerization of functional monomers and cross-
linking agents, or by chemical modification of polymers. Many
heterogeneous membranes have been prepared both by melting and
pressing a mixture of ion-exchange resin and nonfunctional polymer,
or by dissolving or dispersing both functional and support resins in a
solvent and casting into a membrane. Microheterogeneous mem-
branes have been made by block and graft polymerization.
No membrane and no set of membrane properties has universal
applicability. Manufacturers who service multiple applications have a
variety of commercial membranes. One firm lists 20 different mem-
branes having a broad spectrum of properties.
Cation-Exchange Membranes Polystyrene copolymerized with
divinylbenzene, then sulfonated, is the major building block for
cation-exchange membranes. These membranes have reasonable sta-
bility and versatility and are highly ionized over most of the pH range.
Other chemistries mentioned in the literature include carboxylic acid
membranes based on acrylic acid, PO3
, and SeO3
Many specialty membranes have been produced for electrodialysis
applications. A notable example is a membrane selective to monova-
lent cations made by placing a thin coating of positive charge on the
cation-exchange membrane. Charge repulsion for polyvalent ions is
much higher than that for monovalent ions, but the resistance of the
membrane is also higher.
Anion-Exchange Membranes Quaternary amines are the major
charged groups in anion-exchange membranes. Polystyrene-divinyl-
benzene polymers are common carriers for the quaternary amines.
The literature mentions other positive groups based on N, P, and S.
Anion-exchange membranes are problematic, for the best cations are
less robust chemically than their cation exchange counterparts. Since
most natural foulants are colloidal polyanions, they adhere preferen-
tially to the anion-exchange membrane, and since the anion-exchange
membrane is exposed to higher local pH there is a greater likelihood
that precipitates will form there.
Membrane Efficiency The permselectivity of an ion-exchange
membrane is the ratio of the transport of electric charge through the
membrane by specific ions to the total transport of electrons. Mem-
branes are not strictly semipermeable, for coions are not completely
excluded, particularly at higher feed concentrations. For example, the
Donnan equilibrium for a univalent salt in dilute solution is:
Co is the concentration of coions (the ions having the same
electrical charge as the fixed charges on the membrane); CCo
concentration of such coions in the ambient solution; CR
is the con-
centration of fixed charges in the gel water of the membrane; and
and γ Ϯ
are, respectively, the geometric mean of the activity coeffi-
cients of the salt ions in the ambient solution and in the membrane.
Equation (20-101) is applicable only when CCo
. Since the
membrane properties are constant, coion transport rises roughly with
the square of concentration.
Process Description Figure 20-81 gives a schematic view of an
ED cell pair, showing the salt-concentration profile. In the solution
compartment on the left, labeled “Diluate,” anions are being attracted
to the right by the anode. The high density of fixed cations in mem-
brane “A” is balanced by a high mobile anion concentration within it.
Cations move toward the cathode at the left, and there is a similar high
mobile cation concentration in the fixed anion membrane “C” sepa-
rating the depleting compartment from the concentrating compart-
ment further left (not shown). For the anion permeable membrane
“A,” two boundary layers are shown. They represent depletion in the
boundary layer on the left, and an excess in the boundary layer on the
right, both due to the concentration polarization effects common to all
membrane processes—ions must diffuse through a boundary layer
whether they are entering or leaving a membrane. That step must pro-
ceed down a concentration gradient.
With every change in ion concentration, there is an electrical effect
generated by an electrochemical cell. The anion membrane shown in
the middle has three cells associated with it, two caused by the con-
centration differences in the boundary layers, and one resulting from
the concentration difference across the membrane. In addition, there
are ohmic resistances for each step, resulting from the E/I resistance
through the solution, boundary layers, and the membrane. In solution,
current is carried by ions, and their movement produces a friction
effect manifested as a resistance. In practical applications, I2
are more important than the power required to move ions to a com-
partment with a higher concentration.
Transfer of Ions Mass transfer of ions in ED is described by
many electrochemical equations. The equations used in practice are
empirical. If temperature, the flux of individual components, elec-
troosmotic effects, streaming potential and other indirect effects are
MEMBRANE SEPARATION PROCESSES 20-67
FIG. 20-80 Schematic diagram of a cation-exchange membrane showing the
polymer matrix with fixed negative charges and mobile positive counterions.
The density of fixed negative charges is sufficient to prevent the passage
(exchange) of anions.
minor, an equation good to a reasonable level of approximation is:
Jn = CmnUmn ∆ϕ/∆x (20-102)
where J is the component flux through the membrane kmol/m2
is the concentration in phase m of component n, kmol/m3
, Umn is the
ion mobility of n in m, m2
/v⋅s, ϕ is the electrical potential, volts, and x
is distance, m.
Equation (20-102) assumes that all mass transport is caused by an
electrical potential difference acting only on cations and anions.
Assuming the transfer of electrical charges is due to the transfer of
i = F Αn
|zn Jn| (20-103)
where i is the current density, amperes/m2
, F is the Faraday constant
and z is the valence. The transport number, T, is the ratio of the cur-
rent carried by an ion to the current carried by all ions.
Tn = (20-104)
The transport number is a measure of the permselectivity of a mem-
brane. If, for example, a membrane is devoid of coions, then all cur-
rent through the membrane is carried by the counterion, and the
transport number = 1. The transport numbers for the membrane and
the solution are different in practical ED applications.
Concentration Polarization As is shown in the flow schematic,
ions are depleted on one side of the membrane and enriched on the
other. The ions leaving a membrane diffuse through a boundary layer
into the concentrate, so the concentration of ions will be higher at
the membrane surface, while the ions entering a membrane diffuse
through a boundary layer from the diluate, so the bulk concentration
in the diluate must be higher than it is at the membrane. These effects
occur because the transport number of counterions in the membrane
is always very much higher than their transport number in the
solution. Were the transport numbers the same, the boundary layer
effects would vanish. This concentration polarization is similar to that
experienced in reverse osmosis, except that it has both depletion and
enrichment components. The equations governing concentration
polarization and depolarization of a membrane are given in the sec-
tion describing ultrafiltration. The depolarizing strategies used for
ED are similar to those employed in other membrane processes, as
they involve induced flow past the membrane.
Two basic flow schemes are used: tortuous path flow and sheet flow
(Fig. 20-82). Tortuous path spacers are cut to provide a long path
between inlet and outlet, providing a relatively long residence time
and high velocity past the membrane. The flow channel is open. Sheet
flow units have a net spacer separating the membranes. Mass transfer
is enhanced either by the spacer or by higher velocity.
Enhanced depolarization requires capital equipment and energy,
but it achieves savings in overall capital costs (permits the use of a
smaller stack) and energy (permits lower voltage.) The designer’s task
is to achieve an optimum balance in these requirements. Sheet-flow
units have lower capital and operating costs in general, yet both sheet-
flow and tortuous-flow units remain competitive. The fact that most
ED reversal units (see below) are tortuous flow, and that ED reversal
is the dominant technology for water and many waste treatment appli-
cations may explain the paradox.
Limiting Current Density As the concentration in the diluate
becomes ever smaller, or as the current driving the process is in-
creased, eventually a situation arises in which the concentration of ions
at the membranes surrounding the diluate compartment approaches
zero. When that occurs, there are insufficient ions to carry additional
current, and the cell has reached the limiting current. Forcing the
voltage higher results in the dissociation of water at the membrane,
giving rise to a dramatic increase in pH due to OH−
20-68 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-81 Concentration profile of electrolyte across an operating ED cell. Ion passage through the
membrane is much faster than in solution, so ions are enriched or depleted at the cell-solution interface.
“d” is the concentration boundary layer. The cell gap ∆z should be small. The ion concentration in the
membrane proper will be much higher than shown. (Courtesy Elsevier.)
FIG. 20-82 Schematic of two ways to pass solution across an ED membrane.
Tortuous flow (left) uses a special spacer to force the solution through a narrow,
winding path, raising its velocity, mass transfer, and pressure drop. Sheet feed
(right) passes the solution across the plate uniformly, with lower pressure drop
and mass transfer. (Courtesy Elsevier.)
from the anion-exchange membrane. The high pH promotes precipi-
tation of metal hydroxides and CaCO3 on the membrane surface lead-
ing to flow restrictions, poor mass transfer, and subsequent membrane
damage. Once a precipitate forms, its presence initiates a vicious cycle
fostering the formation of more precipitates. The general expression
for the limiting current density is ilim ∼ (Cbulk diluate) (diluate velocity)m
where m is an experimental constant (often 0.6). Thus, the concentra-
tion at the membrane is limiting, and at constant current that is pro-
portional to the bulk concentration and the mass-transfer coefficient.
Flow in the compartment is laminar but mass transfer is enhanced by
the spacer. A normal operating practice is to operate a stack at around
75 percent of the limiting current.
Process Configuration Figure 20-79 shows a basic cell pair. A
stack is an assembly of many cell pairs, electrodes, gaskets, and mani-
folds needed to supply them. An exploded schematic of a portion of a
sheet-flow stack is shown in Fig. 20-83.
Gaskets are very important, since they not only keep the streams
separated and prevent leaks from the cell, they have the manifolds to
conduct feeds, both concentrate and diluate, built into them. No other
practical means of feeding the stack is used in the very cramped space
required by the need to keep cells thin because the diluate has very
low conductivity. The manifolds are formed by aligning holes in mem-
brane and gasket.
The membranes are supported and kept apart by feed spacers. A
typical cell gap is 0.5–2 mm. The spacer also helps control solution
distribution and enhances mass transfer to the membrane. Given that
an industrial stack may have up to 500 cell pairs, assuring uniform flow
distribution is a major design requirement.
Since electrodialysis membranes are subject to fouling, it is some-
times necessary to disassemble a stack for cleaning. Ease of reassem-
bly is a feature of ED.
Process Flow The schematic in Fig. 20-79 may imply that the feed
rates to the concentrate and diluate compartments are equal. If they
are, and the diluate is essentially desalted, the concentrate would leave
the process with twice the salt concentration of the feed. A higher ratio
is usually desired, so the flow rates of feed for concentrate and feed for
diluate can be independently controlled. Since sharply differing flow
rates lead to pressure imbalances within the stack, the usual procedure
is to recirculate the brine stream using a feed-and-bleed technique.
This is usually true for ED reversal plants. Some nonreversal plants use
slow flow on the brine side avoiding the recirculating pumps. Diluate
production rates are often 10ϫ brine-production rates.
Electrodes No matter how many cells are put in series, there will
be electrodes. The more cells, the less the relative importance of the
electrodes. The cathode reactions are relatively mild, and depending
on the pH:
2H2O + 2e−
Anode reactions can be problematic. The anode may dissolve or be
oxidized. Or, depending on the pH and the chloride concentration:
2H2O → O2 + 4H+
→ O2 + 2H2O + 4e−
→ Cl2 + 2e−
Dissolution of metal is avoided by selecting a resistant material such as
Pt, Pt coated on Ti, or Pt on Nb. Base metals are sometimes used, as
are graphite electrodes.
Electrode isolation is practiced to minimize chlorine production
and to reduce fouling. A flush solution free of chlorides or with
reduced pH is used to bathe the electrodes in some plants. Further
information on electrodes may be found in a work by David [“Elec-
trodialysis,” pp. 496–499, in Porter (ed.), op. cit.].
Peripheral Components In addition to the stack, a power supply,
pumps for diluate and concentrate, instrumentation, tanks for cleaning,
and other peripherals are required. Safety devices are mandatory given
the dangers posed by electricity, hydrogen, and chlorine.
Pretreatment Feed water is pretreated to remove gross objects
that could plug the stack. Additives that inhibit the formation of scale,
frequently acid, may be introduced into the feed.
Electrodialysis Reversal Two basic operating modes for ED are
used in large-scale installations. Unidirectional operation is the mode
described above in the general explanation of the process. The elec-
trodes maintain their polarity and the ions always move in a constant
direction. ED reversal is an intermittent process in which the polarity
in the stack is reversed periodically. The interval may be from several
minutes to several hours. When the polarity is reversed, the identity of
compartments is also reversed, and diluate compartments become con-
centrate compartments and vice versa. The scheme requires instru-
ments and valves to redirect flows appropriately after a reversal. The
advantages that often justify the cost are a major reduction in mem-
brane scaling and fouling, a reduction in feed additives required to pre-
vent scaling, and less frequent stack-cleaning requirements.
Water Splitting A modified electrodialysis arrangement is used
as a means of regenerating an acid and a base from a corresponding
salt. For instance, NaCl may be used to produce NaOH and HCl.
Water splitting is a viable alternative to disposal where a salt is pro-
duced by neutralization of an acid or base. Other potential applica-
tions include the recovery of organic acids from their salts and the
treating of effluents from stack gas scrubbers. The new component
required is a bipolar membrane, a membrane that splits water into H+
. At its simplest, a bipolar membrane may be prepared by
laminating a cation and an anion membrane. In the absence of mobile
ions, water sorbed in the membrane splits into its components when a
sufficient electrical gradient is applied. The intimate contact of the
two membranes minimizes the problem of the low ionic conductivity
of ion-depleted water. As the water is split, replacement water readily
MEMBRANE SEPARATION PROCESSES 20-69
FIG. 20-83 Exploded view of a sheet-feed ED stack. Manifolds are built into the membranes and
spacers as the practical way to maintain a narrow cell gap. (Courtesy Elsevier.)
diffuses from the surrounding solution. Properly configured, the
process is energy efficient.
A schematic of the production of acid and base by electrodialytic
water dissociation is shown in Fig. 20-84. The bipolar membrane is
inserted in the ED stack as shown. Salt is fed into the center compart-
ment, and base and acid are produced in the adjacent compartments.
The bipolar membrane is placed so that the cations are paired with
ions and the anions are paired with H+
. Neither salt ion penetrates
the bipolar membrane. As is true with conventional electrodialysis,
many cells may be stacked between the anode and the cathode.
If recovery of both acid and base is unnecessary, one membrane is
left out. For example, in the recovery of a weak acid from its salt, the
anion-exchange membrane may be omitted. The process limitations
relate to the efficiency of the membranes, and to the propensity for H+
to migrate through membranes of like fixed charge, limiting
the attainable concentrations of acid and base to 3–5 N. The problem
is at its worst for HCl and least troublesome for organic acids. Ion
leakage limits the quality of the products, and the regenerated acids
and bases are not of high enough quality to use in regenerating a
mixed-bed ion-exchange resin.
Diffusion Dialysis The propensity of H+
membranes is useful in diffusion dialysis. An anion-exchange mem-
brane will block the passage of metal cations while passing hydrogen
ions. This process uses special ion-exchange membranes, but does not
employ an applied electric current.
As an example, in the aircraft industry heavy-aluminum sections are
shaped as airfoils, then masked. The areas where the metal is not
required to be strong are then unmasked and exposed to NaOH to
etch away unneeded metal for weight reduction. Sodium aluminate is
generated, a potential waste problem. Cation-exchange membranes
by a poorly understood mechanism that is not simply the
transport of OH−
with its waters of hydration. The aluminate anion is
retained in the feed stream while the caustic values pass through.
NaOH recovery is high, because all the Na+
participates in the driving
force. There is considerable passage of water due to the osmotic pres-
sure difference as well. This scheme operates efficiently only because
aluminum hydroxide forms highly supersaturated solutions. Hydrox-
ide precipitation within the apparatus is reported to be a minor prob-
lem. Al(OH)3 is precipitated in a downstream crystallizer, and is
reported to be of high quality.
Donnan Dialysis Another nonelectrical process using ED mem-
branes is used to exchange ions between two solutions. The common
application is to use H+
to drive a cation from a dilute compartment to
a concentrated one. A schematic is shown in Fig. 20-85. In the right
compartment, the pH is 0, thus the H+
concentration is 107
than in the pH 7 compartment on the left. H+
diffuses leftward, creat-
ing an electrical imbalance that can only be satisfied by a cation dif-
fusing rightward through the cation-selective membrane. By this
can be “pumped” from left to right against a significant
Electrodialysis-Moderated Ion Exchange The production of
ultrapure water is facilitated by incorporating a mixed-bed ion-
exchange resin between the membranes of an ion-exchange stack.
Already pure water is passed through the bed, while an electric cur-
rent is passed through the stack. Provided the ion-exchange beads are
in contact with each other and with the membranes, an electrical cur-
rent can pass through the bed even though the conductivity of the
very pure water is quite low. In passing, the current conducts any ions
present into adjacent compartments, simultaneously and continuously
regenerating the resin in situ.
Energy Requirements The thermodynamic limit on energy is
the ideal energy needed to move water from a saline solution to a pure
phase. The theoretical minimum energy is given by:
∆G = RT ln (a/as) (20-105)
where ∆G is the Gibbs free energy required to move one mole of
water from a solution, a is the activity of pure water (ϵ1), and as is the
activity of water in the salt solution. In a solution, the activity of water
is approximately equal to the molar fraction of water in the solution.
So that approximate activity is
as = = 1 + (20-106)
where ns is the number of moles of water in the salt solution, ν is the
number of atoms in the salt molecule (2 for NaCl, 3 for CaCl2) and ss
is the number of moles of salt in the salt solution. The ratio of moles
of salt in the salt solution to the number of moles of water in the salt
solution is a very small number for a dilute solution. This permits
using the approximation ln (1 + x) = x, when x is of the magnitude 0.01,
making this an applicable approximation for saline water. That permits
rewriting Eq. (20-105) as
∆G = νRT(ss /ns) (20-107)
where ∆G is still the free energy required to move one mole of water
from the saline solution to the pure water compartment.
The conditions utilized in the above development of minimum
energy are not sufficient to describe electrodialysis. In addition to the
desalination of water, salt is moved from a saline feed to a more con-
centrated compartment. That free-energy change must be added to
the free energy given in Eq. (20-107), which describes the movement
ν ⋅ ss
ns + ν ⋅ ss
20-70 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-84 Electrodialysis water dissociation (water splitting) membrane
inserted into an ED stack. Starting with a salt, the device generates the corre-
sponding acid and base by supplying H+
from the dissociation of water
in a bipolar membrane. (Courtesy Elsevier.)
FIG. 20-85 Schematic of Donnan dialysis using a cation-exchange membrane.
is “pumped” from the lower concentration on the left to a higher concen-
tration on the right maintaining electrical neutrality accompanying the diffusion
from a low pH on the right to a higher pH on the left. The membrane’s
fixed negative charges prevent mobile anions from participating in the process.
of water from salt solution, the reverse of the actions in the diluate
compartment (but having equal free energy). Schaffer & Mintz
develop that change, and after solving the appropriate material bal-
ances, they arrive at a practical simplified equation for a monovalent
ion salt, where activities may be approximated by concentrations:
∆G = RT(Cf − Cd) − ; Cfc = ; Cfd = (20-108)
where Cf is the concentration of ions in the feed, Cd is the concentra-
tion in the diluate, and Cc is the concentration in the concentrate, all
. When Cfc → 1 and Cfd → infinity, the operation is one
approximating the movement of salt from an initial concentration into
an unlimited reservoir of concentrate, while the diluate becomes
pure. This implies that the concentrate remains at a constant salt
concentration. In that case, Eq. (20-108) reduces to RT(Cf − Cd). As a
numerical example of Eq. (20-108) consider the desalting of a feed
with initial concentration 0.05 M to 0.005 M, roughly approximating
the production of drinking water from a saline feed. If 10 ᐉ of product
are produced for every 3 ᐉ of concentrate, the concentrate leaves the
process at 0.2 M. The energy calculated from Eq. (20-108) is 0.067
at 25°C. If the concentrate flow is infinite, Cc = 0.05 M, and
the energy decreases to 0.031 kWh/m3
This minimum energy is that required to move ions only, and that
energy will be proportional to the ionic concentration in the feed. It
assumes that all resistances are zero, and that there is no polarization.
In a real stack, there are several other important energy dissipaters.
One is overcoming the electrical resistances in the many components.
Another is the energy needed to pump solution through the stack to
reduce polarization and to remove products. Either pumping or
desalting energy may be dominant in a working stack.
Energy Not Transporting Ions Not all current flowing in an
electrodialysis stack is the result of the transport of the intended ions.
Current paths that may be insignificant, minor, or significant include
electrical leakage through the brine manifolds and gaskets, and trans-
port of co-ions through a membrane. A related indirect loss of current
is water transport through a membrane either by osmosis or with sol-
vated ions, representing a loss of product, thus requiring increased
Pump Energy Requirements If there is no forced convection
within the cells, the polarization limits the current density to a very
uneconomic level. Conversely, if the circulation rate is too high, the
Cfd − 1
Cfc − 1
energy inputs to the pumps will dominate the energy consumption of
the process. Furthermore, supplying mechanical energy to the cells
raises the pressure in the cells, and raises the pressure imbalance
between portions of the stack, thus the requirements of the confining
gear and the gaskets. Also, cell plumbing is a design problem made
more difficult by high circulation rates.
A rule of thumb for a modern ED stack is that the pumping energy
is roughly 0.5 kWh/m3
, about the same as is required to remove
1700 mg/ᐉ dissolved salts.
Equipment and Economics A very large electrodialysis plant
would produce 500 ᐉ/s of desalted water. A rather typical plant was
built in 1993 to process 4700 m3
/day (54.4 ᐉ/s). Capital costs for this
plant, running on low-salinity brackish feed were $1,210,000 for all
the process equipment, including pumps, membranes, instrumenta-
tion, and so on. Building and site preparation cost an additional
$600,000. The building footprint is 300 m2
. For plants above a thresh-
old level of about 40 m3
/day, process-equipment costs usually scale at
around the 0.7 power, not too different from other process equip-
ment. On this basis, process equipment (excluding the building) for a
/day plant would have a 1993 predicted cost of $665,000.
The greatest operating-cost component, and the most highly vari-
able, is the charge to amortize the capital. Many industrial firms use
capital charges in excess of 30 percent. Some municipalities assign
long amortization periods and low-interest rates, reflecting their cost
of capital. Including buildings and site preparation, the range of capi-
tal charges assignable to 1000 m3
of product is $90 to $350.
On the basis of 1000 m3
of product water, the operating cost ele-
ments (as shown in Table 20-32) are anticipated to be:
SELECTION OF BIOCHEMICAL SEPARATION PROCESSES 20-71
TABLE 20-32 Electrodialysis Operating Costs
$ 66 Membrane-replacement cost (assuming seven-year life)
32 Plant power
16 Filters and pretreatment chemicals
SELECTION OF BIOCHEMICAL SEPARATION PROCESSES
GENERAL REFERENCES: Ahuja (ed.), Handbook of Bioseparations, Academic
Press, London, 2000. Albertsson, Partition of Cell Particles and Macromole-
cules, 3d ed., Wiley, New York, 1986. Belter, Cussler, and Hu, Bioseparations,
Wiley Interscience, New York, 1988. Cooney and Humphrey (eds.), Compre-
hensive Biotechnology, vol. 2, Pergamon, Oxford, 1985. Flickinger and Drew
(eds.), The Encyclopedia of Bioprocess Technology: Fermentation, Biocatalysis,
and Bioseparations, Wiley, New York, 1999. Harrison, Todd, Scott, and
Petrides, Bioseparations Science and Engineering, Oxford University Press,
Oxford, 2003. Hatti-Kaul and Mattiasson (eds.), Isolation and Purification of
Proteins, Marcel Dekker, New York, 2003. Janson and Ryden (eds.), Protein
Purification, VCH, New York, 1989. Ladisch, Bioseparations Engineering: Prin-
ciples, Practice, and Economics, Wiley, New York, 2001. Scopes, Protein Purifi-
cation, 2d ed., Springer-Verlag, New York, 1987. Stephanopoulos (ed.),
Biotechnology, 2d ed., vol. 3, VCH, Weinheim, 1993. Subramanian (ed.),
Bioseparation and Bioprocesss—A Handbook, Wiley, New York, 1998.
Zaslavsky, Aqueous Two-Phase Partitioning—Physical Chemistry and Bioana-
lytical Applications, Marcel Dekker, New York, 1995.
The biochemical industry derives its products from two primary
sources. Natural products are yielded by plants, animal tissue, and flu-
ids, and they are obtained via fermentation from bacteria, molds, and
fungi and from mammalian cells. Products can also be obtained by
recombinant methods through the insertion of foreign DNA directly
into the hosting microorganism to allow overproduction of the prod-
uct in this unnatural environment. The range of bioproducts is enor-
mous, and the media in which they are produced are generally
complex and ill-defined, containing many unwanted materials in addi-
tion to the desired product. The product is almost invariably at low
concentration to start with. The goals of downstream processing oper-
ations include removal of these unwanted impurities, bulk-volume
reduction with concomitant concentration of the desired product,
and, for protein products, transfer of the protein to an environment
where it will be stable and active, ready for its intended application.
This always requires a multistage process consisting of multiple-unit
operations. A general strategy for downstream processing of biological
materials and the types of operations that may be used in the different
steps is shown in Fig. 20-86 [see also Ho, in M. R. Ladisch et al. (eds.),
Protein Purification from Molecular Mechanisms to Large-Scale
Processes, ACS Symp. Ser., 427, ACS, Washington, D.C. (1990), pp.
14–34]. Low-molecular-weight products, generally secondary metabo-
lites such as alcohols, carboxylic and amino acids, antibiotics, and vit-
amins, can be recovered by using many of the standard operations
such as liquid-liquid partitioning, adsorption, and ion exchange,
described elsewhere in this handbook. Biofuel molecules also belong
These items are highly site specific. Power cost is low because the
salinity removed by the selected plant is low. The quality of the feed
water, its salinity, turbidity, and concentration of problematic ionic and
fouling solutes, is a major variable in pretreatment and in conversion.
20-72 ALTERNATIVE SEPARATION PROCESSES
Cell Debris Removal
Final purification and
product polishingDenatured products
Whole broth treatment
Salt or Solvent Removal
FIG. 20-86 General stages in downstream processing for protein production indicating representative types of unit operations used
at each stage.
to this category. They are becoming more and more important. Less
expensive and high-throughput unit operations are needed to make a
biorefinery process economically feasible. Proteins require special
attention, however, as they are sufficiently more complex, their func-
tion depending on the integrity of a delicate three-dimensional ter-
tiary structure that can be disrupted if the protein is not handled
correctly. For this reason, this section focuses primarily on protein
Techniques used in bioseparations depend on the nature of the
product (i.e., the unique properties and characteristics which provide
a “handle” for the separation) and on its state (i.e., whether soluble or
insoluble, intra- or extracellular, etc.). All early isolation and recovery
steps remove whole cells, cellular debris, suspended solids, and col-
loidal particles; concentrate the product; and in many cases, achieve
some degree of purification, all the while maintaining high yield. For
intracellular compounds, the initial harvesting of the cells is important
for their concentration prior to release of the product. Following this
phase, a range of purification steps are employed to remove the
remaining impurities and enhance the product purity; this purification
phase, in turn, is followed by polishing steps to remove the last traces
of contaminating components and process-related additions (e.g.,
buffer salts, detergents) and to prepare the product for storage and/or
distribution. The prevention and/or avoidance of contamination is
another important goal of downstream processing. Therapeutic pro-
teins require very high purities that cannot be measured by weight
percentages alone. To avoid potential harmful effects in humans, the
levels of pyrogens and microbial and viral contaminates in final prod-
ucts must meet stringent safety and regulatory requirements.
Even for good yields of 80 to 95 percent per step, the overall yield
can be poor for any process that requires a large number of steps.
Thus, careful consideration must be given to optimization of the
process in terms of both the unit operations themselves and their
sequencing. A low-yield step should be replaced if improvement is
unattainable to eliminate its impact on the overall yield. It is usually
desirable to reduce the process volume early in the downstream pro-
cessing, and to remove any components that can be removed fairly
easily (particulates, small solutes, large aggregates, nucleic acids, etc.),
so as not to overly burden the more refined separation processes
downstream. Possible shear and temperature damage, and deactiva-
tion by endogenous proteases, must be considered in the selection of
separation processes. Protein stability in downstream processing was
discussed in depth by Hejnaes et al. [in Subramanian (ed.), vol. 2, op.
cit., pp. 31–66].
INITIAL PRODUCT HARVEST AND CONCENTRATION
The initial processing steps are determined to a large extent by the
location of the product species, and they generally consist of cell/broth
separation and/or cell debris removal. For products retained within
the biomass during production, it is first necessary to concentrate the
cell suspension before homogenization or chemical treatment to
release the product. Clarification to remove the suspended solids is
the process goal at this stage.
Regardless of the location of the protein and its state, cell separa-
tion needs to be inexpensive, simple, and reliable, as large amounts of
fermentation-broth dilute in the desired product may be handled. The
objectives are to obtain a well-clarified supernatant and solids of max-
imum dryness, avoiding contamination by using a contained opera-
tion. Centrifugation or crossflow filtration is typically used for cell
separation, and both unit operations can be run in a continuous-flow
mode [Datar and Rosen, in Stephanopoulos (ed.), op. cit., pp.
369–503]. In recent years, expanded-bed adsorption has become an
alternative. It combines broth clarification and adsorption separation
in a single step.
Intracellular products can be present either as folded, soluble pro-
teins or as dense masses of unfolded protein (inclusion bodies). For
these products, it is first necessary to concentrate the cell suspension
before effecting release of the product. Filtration can result in a sus-
pension of cells that can be of any desired concentration up to 15 to 17
percent and that can be diafiltered into the desired buffer system. In
contrast, the cell slurry that results from centrifugation will be that of
either a dry mass (requiring resuspension but substantially free of
residual broth, i.e., from a tubular bowl centrifuge) or a wet slurry
(containing measurable residual broth and requiring additional resus-
pension). During the separation, conditions that result in cell lysis
(such as extremes in temperature) must be avoided. In addition, while
soluble protein is generally protected from shear and external prote-
olysis, these proteins are still subject to thermal denaturation.
Cell Disruption Intracellular protein products are present as
either soluble, folded proteins or inclusion bodies. Intracellular pro-
tein products are very common because Escherichia coli is a main
workhorse for recombinant proteins. E. coli is a gram-negative bac-
terium that precipitates recombinant proteins in the form of inclusion
bodies. Release of folded proteins must be carefully considered.
Active proteins are subject to deactivation and denaturation and thus
require the use of “gentle” conditions. In addition, due consideration
must be given to the suspending medium; lysis buffers are often opti-
mized to promote protein stability and protect the protein from pro-
teolysis and deactivation. Inclusion bodies, in contrast, are protected
by virtue of the protein agglomeration. More stressful conditions are
typically employed for their release, which includes going to higher
temperatures if necessary. For “native” proteins, gentler methods and
temperature control are required.
The release of intracellular protein product is achieved through
rupture of the cell walls, and release of the protein product to the sur-
rounding medium, through either mechanical or nonmechanical
means, or through chemical, physical, or enzymatic lysis [Engler, in
Cooney and Humphrey (eds.), op. cit., pp. 305–324; Schutte and
Kula, in Stephanopoulos (ed.), op. cit., pp. 505–526]. Mechanical
methods use pressure, as in the Manton/APV-Gaulin/French Press, or
the Microfluidizer, or mechanical grinding, as in ball mills, the latter
being used typically for flocs and usually only for natural products.
Nonmechanical means include use of desiccants or solvents, while cell
lysis can also be achieved through physical means (osmotic shock,
freeze/thaw cycles), chemical (detergents, chaotropes), or enzymatic
In a high-pressure homogenizer, a pressurized cell suspension is
forced through a valve and undergoes a rapid pressure change from
up to 50 MPa to the atmospheric pressure. This results in the instant
rupture of cells. Product release, which generally follows first-order
kinetics, occurs through impingement of the high-velocity cell sus-
pension jet on the stationary surfaces, and possibly also by the high-
shear forces generated during the acceleration of the liquid through
the gap. While sufficiently high pressures can be attained using com-
mercially available equipment to ensure good release in a single pass,
the associated adiabatic temperature increases (∼1.8°C/1000 psig)
may cause unacceptable activity losses for heat-labile proteins. Fur-
ther denaturation can occur on exposure to the lysis medium. Thus,
multiple passes may be preferred, with rapid chilling of the processed
cell suspension between passes. The number of passes and the heat
removal ability should be carefully optimized. The efficiency of the
process depends on the homogenizing pressure and the choice of the
valve unit, for which there are many designs available. Materials of
construction are important to minimize erosion of the valve, to pro-
vide surface resistance to aggressive cleaning agents and disinfectants,
and to permit steam cleaning and sanitization.
The release of inclusion bodies, in contrast, may follow a different
strategy. Since inclusion bodies are typically recovered by centrifuga-
tion, it is often advantageous to send the lysate through the homoge-
nizer with multiple passes to decrease the particle size of the cell
debris. Since the inclusion bodies are much denser than the cell
debris, the debris, now much reduced in size, can be easily separated
from the inclusion bodies by centrifugation at low speeds. The inclu-
sion bodies may be resuspended and centrifuged multiple times
(often in the presence of low concentrations of denaturants) to clean
up these aggregates. Since the inclusion bodies are already denatured,
temperature control is not as important as in the case of native pro-
Another popular method for cell disruption is to use a bead mill. In a
bead mill, a cell suspension is mixed with glass or metal beads and agi-
tated by using a rotating agitator at high speed. Bead mills have a con-
trollable residence time compared with high-pressure homogenization.
SELECTION OF BIOCHEMICAL SEPARATION PROCESSES 20-73
However, they are susceptible to channeling and also fracturing of the
beads. Tough cells require multiple passes to achieve a desired yield.
Chemical lysis, or solubilization of the cell wall, is typically carried
out by using detergents such as Triton X-100, or the chaotropes urea,
and guanidine hydrochloride. This approach does have the disadvan-
tage that it can lead to some denaturation or degradation of the prod-
uct. While favored for laboratory cell disruption, these methods are not
typically used at the larger scales. Enzymatic destruction of the cell
walls is also possible, and as more economical routes to the develop-
ment of appropriate enzymes are developed, this approach could find
industrial application. Again, the removal of these additives is an issue.
Physical methods such as osmotic shock, in which the cells are
exposed to high salt concentrations to generate an osmotic pressure
difference across the membrane, can lead to cell wall disruption. Sim-
ilar disruption can be obtained by subjecting the cells to freeze/thaw
cycles, or by pressurizing the cells with an inert gas (e.g., nitrogen) fol-
lowed by a rapid depressurization. These methods are not typically
used for large-scale operations.
On homogenization, the lysate may drastically increase in viscosity
due to DNA release. This can be ameliorated to some extent by using
multiple passes to reduce the viscosity. Alternatively, precipitants or
nucleic acid digesting enzymes can be used to remove these viscosity-
For postlysis processing, careful optimization must be carried out
with respect to pH and ionic strength. Often it is necessary to do a
buffer exchange. Cell debris can act as an ion exchanger and bind pro-
teins ionically, thus not allowing them to pass through a filtration
device or causing them to be spun out in a centrifuge. Once optimal
conditions are found, these conditions can be incorporated in the lysis
buffer by either direct addition (if starting from cell paste) or diafil-
tration (if starting from a cell concentrate).
Protein Refolding Although protein refolding is not a biosepa-
ration operation, it is an integral part of a downstream process for the
production of an inactive (typically intracellular) protein. So far, it still
remains a challenging art casting an uncertainty on the success of a
process. Early in the product development stage, animal cell culture
may be required to produce many bioactive candidates for initial
screening of efficacy. To commercialize the product protein, it may be
reexpressed in a far more economic and productive host such as E.
coli. The commercial products of recombinant DNA technology are
frequently not produced in their native, biologically active form,
because the foreign hosts such as E. coli in which they are produced
lack the appropriate apparatus for the folding of the proteins. Thus,
the overproduced proteins are generally recovered as refractile or
inclusion bodies, or aggregates, typically 1 to 3 µm in size, and all cys-
teine residues are fully reduced. It is necessary at some stage in the
processing to dissolve the aggregates and then refold them to obtain
the desired biologically active product [Cleland and Wang, in
Stephanopoulos (ed.), op. cit., pp. 527–555].
Advantages of inclusion bodies in the production stage are their
ease of separation by centrifugation following cell disruption, because
of their size and density, and their provision of excellent initial-
purification possibilities, as long as impurities are not copurified to
any significant extent with the inclusion bodies. They also provide a
high expression level and prevent endogenous proteolysis. There can
be, however, significant product loss during protein refolding to the
active form. Figure 20-87 shows a typical process for refolding. The
inclusion body which is released from the host cell by cell disruption
is washed and then solubilized by using a denaturant such as guani-
dine hydrochloride (4 to 9 M), urea (7 to 8 M), sodium thiocyanate
(4 to 9 M), or detergents such as Triton X-100 or sodium dodecyl sul-
fate. This step disrupts the hydrogen and ionic bonds to obtain fully
denatured and stretched peptide chains. For proteins with disulfide
bonds, addition of appropriate reducing agents (e.g., beta mercap-
toethanol) is required to break all incorrectly formed intramolecular
disulfide bonds. To permit proper refolding of the protein, it is neces-
sary to remove the denaturant or detergent molecules from the sur-
roundings of the stretched and solubilized peptides. This will initiate
self-refolding of the protein molecules. For proteins with disulfide
bonds, an oxidative reaction with oxygen or other oxidants is required
to join two free SH groups to form an SᎏS covalent bond. This in
vitro refolding operation is traditionally achieved by dilution with a
refolding buffer. However, misfolding or aggregation is usually found
in the refolding process. Analysis of in vitro refolding kinetics shows
that there is at least an intermediate (I) between the unfolded protein
(U) and the fully active refolded native protein (N), as illustrated
below [Kuwajima and Arai, in R. H. Pain (ed.), Mechanism of Protein
Folding, 2d ed., Oxford University Press, Oxford, 2000, pp. 138–171;
Tsumoto et al., Protein Expr. Purif., 28, 1–8 (2003)].
20-74 ALTERNATIVE SEPARATION PROCESSES
U I N
FIG. 20-87 Illustration of a refolding process for a protein from inclusion body.
The process to convert U to I may be fast. The process for I to N,
however, may be slow and highly reversible. Some intermediate mol-
ecules may form aggregates (A) or misfolded proteins (M). Figure 20-
87 is a simplified refolding pathway. In reality, the situation can be
more complicated where several intermediates (I1, I2, I3, etc.) are
present with numerous possibilities of aggregation and misfolding.
For most proteins, refolding is a self-assembly process that follows a
first-order kinetics. Aggregation, on the other hand, involves interac-
tions between two or more molecules and follows the second- or
higher-order kinetics. Therefore, in vitro refolding at higher protein
concentrations would lead to the formation of more aggregates. Many
observations have shown that a low final protein concentration, usu-
ally 10 to 100 µg/mL, is required for dilution refolding [Schlegl et al.,
Chem. Engng. Sci., 60, 5770–5780 (2005)]. In an industrial process,
this strategy generally features large volumes of buffers, exerting an
extra burden for subsequent purification steps because the concentra-
tion is low. Optimization of dilution strategy, such as the way of dilu-
tion, the speed of dilution, and the solution composition of a refolding
buffer, is beneficial for an increased refolding yield.
A number of studies [Kuwajima and Arai, loc. cit.; Sadana, Biotech.
Bioeng., 48, 481–489 (1995)] demonstrated that the presence of some
molecules in the refolding buffer may suppress misfolding or aggrega-
tion. Molecular chaperones such as GroES and GroEL can promote
correct refolding, but they are very expensive. Other molecules have
been tried as artificial chaperones. The most commonly used mole-
cules are polyethylene glycol (PEG) of various molecular weights and
concentrations, L-arginine (0.4 to 1 M), low concentrations of denatu-
rants such as urea (1 to 2 M) and guanidine hydrochloride (0.5 to 1.5 M),
and some detergents such as SDS, CTAB, and Triton X-100. The effects
of these additives on protein refolding are still under investigation. A
new direction of protein refolding involves the use of chromatographic
techniques [Li et al., Protein Expr. Purif., 33, 1–10 (2004)]. Size exclu-
sion, ion exchange, hydrophobic interaction, and metal chelate affinity
chromatographic techniques have been studied with successful results.
Chromatographic refolding explores the interaction between the pro-
tein molecule to be refolded and the packing medium in the column. It
may reduce the interaction between protein molecules and increase the
chance of self-assembly with the aid of the functional groups and pores
in the matrix of the packing medium. An advantage for chromato-
graphic refolding is the availability of gradient elution that creates a
gradual change of the solution environment, leading to a gentle removal
of the denaturant and a gradual change of favorable conditions such as
pH and the artificial chaperone concentration. Simultaneous refolding
and partial purification are possible with this new technique.
For extracellular products, which are invariably water-soluble, the
first step is the removal of cells and cell debris using a clarification
method and, in the case of typical protein products, the removal of
dissolved low-molecular-weight compounds. This must be done under
relatively gentle conditions to avoid undesired denaturation of the
product. Again, either filtration or centrifugation can be applied. Fil-
tration results in a cell-free supernatant with dilution associated with
the diafiltration of the final cell slurry. Centrifugation, regardless of
the mode, will result in a small amount of cells in the centrate, but
there is no dilution of the supernatant. During the process develop-
ment careful studies should be conducted to examine the effects of
pH and ionic strength on the yield, as cells and cell debris may retain
the product through charge interactions. If the broth or cell morphol-
ogy does not allow for filtration or if dry cell mass is required, tubular-
bowl centrifugation is typically utilized. Note that plant and animal
cells cannot sustain the same degree of applied shear as can microbial
cells, and thus crossflow filtration or classical centrifugation may not
be applicable. Alternatives using low-shear equipment under gentle
conditions are often employed in these situations.
For whole broths the range of densities and viscosities encountered
affects the concentration factor that can be attained in the process,
and can also render crossflow filtration uneconomical because of the
high pumping costs, and so on. Often, the separation characteristics of
the broth can be improved by broth conditioning using physicochem-
ical or biological techniques, usually of a proprietary nature. The
important characteristics of the broth are rheology and conditioning.
Clarification Using Centrifugation Centrifugation relies on
the enhanced sedimentation of particles of density different from that
of the surrounding medium when subjected to a centrifugal force field
[Axelsson, in Cooney and Humphrey (eds.), op. cit., pp. 325–346] (see
also Sec. 18, “Liquid-Solid Operations and Equipment”). Advantages
of centrifugal separations are that they can be carried out continuously
and have short retention times, from a fraction of a second to seconds,
which limit the exposure time of sensitive biologicals to shear stresses.
Yields are high, provided that temperature and other process condi-
tions are adequately controlled. They have small space requirements,
and an adjustable separation efficiency makes them a versatile unit
operation. They can be completely closed to avoid contamination, and,
in contrast to filtration, no chemical external aids are required that can
contaminate the final product. The ability now to contain the aerosols
typically generated by centrifuges adds to their operability and safety.
Sedimentation rates must be sufficiently high to permit separation,
and they can be enhanced by modifying solution conditions to pro-
mote the aggregation of proteins or impurities. An increase in precip-
itation of the contaminating species can often be accomplished by a
reduction in pH or an elevation in temperature. Flocculating agents,
which include polyelectrolytes, polyvalent cations, and inorganic salts,
can cause a 2000-fold increase in sedimentation rates. Some examples
are polyethylene imine, EDTA, and calcium salts. Cationic biopro-
cessing aids (cellulosic or polymeric) reduce pyrogen, nucleic acid,
and acidic protein loads which can foul chromatography columns. The
removal of these additives during both centrifugation and subsequent
processing must be clearly demonstrated.
There are many different types of centrifuges, classified according to
the way in which the transport of the sediment is handled [Medronho, in
Hatti-Kaul and Mattiasson (eds.), op. cit., pp. 131–190]. The selection of
a particular centrifuge type is determined by its capacity for handling
sludge; the advantages and disadvantages of various separator types are
discussed by Axelsson [in Cooney and Humphrey (eds.), op. cit., pp.
325–346]. Solids-retaining centrifuges are operated in a semibatch
mode, as they must be shut down periodically to remove the accumu-
lated solids; they are primarily used when solids concentrations are low,
and they have found application during the clarification and simultane-
ous separation of two liquids. In solids-ejecting centrifuges, the solids are
removed intermittently either through radial slots or axially while the
machine is running at full speed. These versatile machines can be used
to handle a variety of feeds, including yeast, bacteria, mycelia, antibiotics,
enzymes, and so on. Solids-discharging nozzle centrifuges have a large
capacity and can accommodate up to 30 percent solids loading. Decanter
centrifuges consist of a drum, partly cylindrical and partly conical, and an
internal screw conveyor for transport of the solids, which are discharged
at the conical end; liquids are discharged at the cylindrical end. Levels
within the drum are set by means of external nozzles.
Continuous-flow units, the scroll decanter and disk-stack cen-
trifuges, are easiest to use from an operational perspective; shutdown
of the centrifuge during the processing of a batch is not expected.
While the disk-stack centrifuge enjoys popularity as a process instru-
ment within the pharmaceutical and biotechnology industries, the
precise timing of solids ejection and the continuous high-speed nature
of the device make for complex equipment and frequent mainte-
nance. It is often used to harvest cells, since the solids generated are
substantially wet and could lead to measurable yield losses in extracel-
lular product systems. For intracellular product processing, the wet
cell sludge is easily resuspended for use in subsequent processing.
The tubular bowl, in contrast, is a semibatch processing unit owing
to the limited solids capacity of the bowl. The use of this unit requires
shutdown of the centrifuge during the processing of the batch. The
semibatch nature of these centrifuges can thus greatly increase pro-
cessing cycle times. The introduction of disposable sheets to act as
bowl liners has significantly impacted turnaround times during pro-
cessing. The dry nature of the solids generated makes the tubular-
bowl centrifuge well suited for extracellular protein processing, since
losses to the cell sludge are minimal. In contrast, the dry, compact
nature of the sludge can make the cells difficult to resuspend. This can
be problematic for intracellular protein processing where cells are
homogenized in easily clogged, mechanical disrupters.
Clarification Using Microfiltration Crossflow filtration
(microfiltration includes crossflow filtration as one mode of operation
in “Membrane Separation Processes,” which appears earlier in this
section) relies on the retention of particles by a membrane. The driv-
ing force for separation is pressure across a semipermeable membrane,
while a tangential flow of the feed stream parallel to the membrane
surface inhibits solids settling on and within the membrane matrix
(Datar and Rosen, loc. cit.).
Microfiltration is used for the removal of suspended particles,
recovery of cells from fermentation broth, and clarification of
homogenates containing cell debris. Particles removed by microfiltra-
tion typically average greater than 500,000 nominal molecular weight
[Tutunjian, in Cooney and Humphrey (eds.), op. cit., pp. 367–381;
Gobler, in Cooney and Humphrey (eds.), op. cit., pp. 351–366]. Ultra-
filtration focuses on the removal of low-molecular-weight solutes and
proteins of various sizes, and it operates in the less than 100,000
nominal-molecular-weight cutoff (NMWCO) range [Le and Howell,
in Cooney and Humphrey (eds.), op. cit., pp. 383–409]. Both opera-
tions consist of a concentration segment (of the larger particles) fol-
lowed by diafiltration of the retentate [Tutunjian, in Cooney and
Humphrey (eds.), op. cit., pp. 411–437].
Generally, the effectiveness of the separation is determined not by
the membrane itself, but rather by the formation of a secondary or
dynamic membrane caused by interactions of the solutes and particles
with the membrane. The buildup of a gel layer on the surface of an
ultrafiltration membrane owing to rejection of macromolecules can
provide the primary separation characteristics of the membrane.
Similarly, with colloidal suspensions, pore blocking and bridging of
SELECTION OF BIOCHEMICAL SEPARATION PROCESSES 20-75
pore entries can modify the membrane performance, while mole-
cules of size similar to the membrane pores can adsorb on the pore
walls, thereby restricting passage of water and smaller solutes. Media
containing poorly defined ingredients may contain suspended solids,
colloidal particles, and gel-like materials that prevent effective microfil-
tration. In contrast to centrifugation, specific interactions can play a sig-
nificant role in membrane separation processes.
The factors to consider in the selection of crossflow filtration include
the flow configuration, tangential linear velocity, transmembrane pres-
sure drop (driving force), separation characteristics of the membrane
(permeability and pore size), size of particulates relative to the mem-
brane pore dimensions, low protein-binding ability, and hydrodynamic
conditions within the flow module. Again, since particle-particle and
particle-membrane interactions are key, broth conditioning (ionic
strength, pH, etc.) may be necessary to optimize performance.
Selection of Cell-Separation Unit Operation The unit opera-
tion selected for cell separations can depend on the subsequent sepa-
ration steps in the train. In particular, when the operation following
cell separation requires cell-free feed (e.g., chromatography), filtra-
tion is used, since centrifugation is not absolute in terms of cell sepa-
ration. In addition, if cells are to be stored (i.e., they contain the
desired product) because later processing is more convenient (e.g.,
only two-shift operation, facility competes for equipment with other
products, batch is too big for single pass in equipment), it is generally
better to store the cells as a frozen concentrate than a paste, since the
concentrate thaws more completely, avoiding small granules of
unfrozen cell solids that can foul homogenizers, columns, and filters.
Here the retentate from filtration is desired, although the wet cell
mass from a disc stack-type centrifuge may be used.
Centrifugation is generally necessary for complex media used to
make natural products, for while the media components may be sifted
prior to use, they can still contain small solids that can easily foul fil-
ters. The medium to be used should be tested on a filter first to deter-
mine the fouling potential. Some types of organisms, such as
filamentous organisms, may sediment too slowly owing to their larger
cross sections, and they are better treated by filtration (mycelia have
the potential to easily foul tangential-flow units; vacuum-drum filtra-
tion using a filter aid, e.g., diatomaceous earth, should also be consid-
ered). Often the separation characteristics of the broth can be
improved by broth conditioning using physicochemical or biological
techniques, usually of a proprietary nature.
Regardless of the machine device, centrifuges are typically mainte-
nance-intensive. Filters can be cheaper in terms of capital and main-
tenance costs and should be considered first unless centrifugal
equipment already exists. Small facilities (< 1000 L) use filtration,
since centrifugation scale-down is constrained by equipment availabil-
ity. Comparative economics of the two classes of operations are dis-
cussed by Datar and Rosen (loc. cit.).
Initial purification is the rough purification (considered by many peo-
ple as isolation) to prepare a feed for subsequent high-resolution
steps. In initial purification steps the goal is to obtain concentration
with partial purification of the product, which is recovered as a pre-
cipitate (precipitation), a solution in a second phase (liquid-liquid par-
titioning), or adsorbed to solids (adsorption, chromatography).
Precipitation Precipitation of products, impurities, or contami-
nants can be induced by the addition of solvents, salts, or polymers to
the solution; by increasing temperature; or by adjusting the solution
pH (Scopes, op. cit., pp. 41–71; Ersson et al., in Janson and Ryden, op.
cit., pp. 3–32). This operation is used most often in the early stages of
the separation sequence, particularly following centrifugation, filtra-
tion, and/or homogenization steps. Precipitation is often carried out in
two stages, the first to remove bulk impurities and the second to pre-
cipitate and concentrate the target protein. Generally, amorphous
precipitates are formed, owing to occlusion of salts or solvents, or to
the presence of impurities.
Salts can be used to precipitate proteins by “salting out” effects.
The effectiveness of various salts is determined by the Hofmeister
series, with anions being effective in the order citrate > PO4
, and cations according to NH4
(Ersson et al., op. cit., p. 10; Belter et al., op. cit., pp. 221–236).
Salts should be inexpensive owing to the large quantities used in pre-
cipitation operations. Ammonium sulfate appears to be the most popular
precipitant because it has an effective cation and an effective anion, high
solubility, easy disposal, and low cost. Drawbacks to this approach
include low selectivity, high sensitivity to operating conditions, and
downstream complications associated with salt removal and disposal of
the high-nitrogen-content stream. Generally, aggregates formed on pre-
cipitation with ammonium sulfate are fragile, and are easily disrupted by
shear. Thus, these precipitation operations are, following addition of salt,
often aged without stirring before being fed to a centrifuge by gravity
feed or using low-shear pumps (e.g., diaphragm pumps).
The organic solvents most commonly used for protein precipitation
are acetone and ethanol (Ersson et al., op. cit.). These solvents can
cause some denaturation of the protein product. Temperatures below
0°C can be used, since the organic solvents depress the freezing point
of the water. The precipitate formed is often an extremely fine powder
that is difficult to centrifuge and handle. With organic solvents, in-line
mixers are preferred, as they minimize solvent-concentration gradi-
ents and regions of high-solvent concentrations, which can lead to sig-
nificant denaturation and local precipitation of undesired components
typically left in the mother liquors. In general, precipitation with
organic solvents at lower temperature increases yield and reduces
denaturation. It is best carried out at ionic strengths of 0.05 to 0.2 M.
Water-soluble polymers and polyelectrolytes (e.g., polyethylene
glycol, polyethylene imine polyacrylic acid) have been used success-
fully in protein precipitations, and there has been some success in
affinity precipitations wherein appropriate ligands attached to poly-
mers can couple with the target proteins to enhance their aggregation.
Protein precipitation can also be achieved by using pH adjustment,
since proteins generally exhibit their lowest solubility at their isoelec-
tric point. Temperature variations at constant salt concentration allow
for fractional precipitation of proteins.
Precipitation is typically carried out in standard cylindrical tanks
with low-shear impellers. If in-line mixing of the precipitating agent is
to be used, this mixing is employed just prior to the material’s entering
the aging tank. Owing to their typically poor filterability, precipitates
are normally collected by using a centrifugal device.
Liquid-Liquid Partitioning Liquid-liquid partitioning (see also
Sec. 15 on liquid-liquid extraction) involving an organic solvent is
commonly known as solvent extraction or extraction. Solvent extrac-
tion is routinely used to separate small biomolecules such as antibi-
otics and amino acids. However, it is typically not suitable for protein
fractionation with only a few exceptions because organic solvents may
cause protein denaturation or degradation. A recent review of solvent
extraction for bioseparations including a discussion on various para-
meters that can be controlled for solvent extraction was given by Gu
[in Ahuja (ed.), op. cit., pp. 365–378]. As a replacement of solvent
extraction, aqueous two-phase partitioning is typically used for protein
purification. It uses two water-soluble polymers (and sometimes with
some salts when polyelectrolytes are involved) to form two aqueous
phases [Albertsson, op. cit.; Kula, in Cooney and Humphrey (eds.),
op. cit., pp. 451–471]. Both phases contain water but differ in polymer
(and salt) concentration(s). The high water content, typically greater
than 75 percent, results in a biocompatible environment not attain-
able with traditional solvent extraction systems. Biomolecules such as
proteins have different solubilities in the two phases, and this provides
a basis for separation. Zaslavsky (op. cit., pp. 503–667) listed 163
aqueous two-phase systems including PEG-dextran-water, PEG-
polyvinylmethylether-water, PEG-salt-water, polyvinylpyrrolidone-
dextran-water, polyvinylalcohol-dextran-water, and Ficoll-dextran-water
systems. Partitioning between the two aqueous phases is controlled by
the polymer molecular weight and concentration, protein net charge
and size, and hydrophobic and electrostatic interactions. Aqueous
two-phase polymer systems are suitable for unclarified broths since
particles tend to collect at the interface between the two phases, mak-
ing their removal very efficient. They can also be used early on in the
processing train for initial bulk-volume reduction and partial purifica-
tion. One of the drawbacks of these systems is the subsequent need
for the removal of phase-forming reagents.
20-76 ALTERNATIVE SEPARATION PROCESSES
Affinity partitioning is carried out by adding affinity ligands to an
aqueous two-phase partitioning system. The biospecific binding of a
biomolecule with the ligand moves the biomolecule to a preferred
phase that enhances the partitioning of the biomolecule [Johansson
and Tjerneld, in Street (ed.), Highly Selective Separations in Biotech-
nology, Blackie Academic & Professional, London, 1994, pp. 55–85].
Diamond and Hsu [in Fiechter (ed.), Advances in Biochemical Engi-
neering/Biotechnology, vol. 47, Springer-Verlag, Berlin-New York,
2002, pp. 89–135] listed several dozens of biomolecules, many of
which are proteins that have been separated by using affinity parti-
tioning. Fatty acids and triazine are the two common types of affinity
ligand while metallated iminodiacetic acid (IDA) derivatives of PEG
such as Cu(II)IDA-PEG can be used for binding with proteins rich in
surface histidines. The drawbacks of affinity partitioning include the
costs of ligands and the need to couple the ligands to the polymers
used in the aqueous two-phase partitioning.
Product recovery from these systems can be accomplished by
changes in either temperature or system composition. Composition
changes can be affected by dilution, backextraction, and micro- and
ultrafiltration. As the value of the product decreases, recovery of the
polymer may take on added significance. A flow diagram showing one
possible configuration for the extraction and product and polymer
recovery operations is shown in Fig. 20-88 [Greve and Kula, J. Chem.
Tech. Biotechnol., 50, 27–42 (1991)]. The phase-forming polymer and
salt are added directly to the fermentation broth. The cells or cell
debris and contaminating proteins report to the salt-rich phase and are
discarded. Following pH adjustment of the polymer-rich phase, more
salt is added to induce formation of a new two-phase system in which
the product is recovered in the salt phase, and the polymer can be recy-
cled. In this example, disk-stack centrifuges are used to enhance the
phase separation rates. Other polymer recycling options include
extraction with a solvent or supercritical fluid, precipitation, or diafil-
tration. Electrodialysis can be used for salt recovery and recycling.
Reversed micellar solutions can also be used for the selective
extraction of proteins [Kelley and Hatton, in Stephanopoulos (ed.),
op. cit., pp. 593–616]. In these systems, detergents soluble in an oil
phase aggregate to stabilize small water droplets having dimensions
similar to those of the proteins to be separated. These droplets can
host hydrophilic species such as proteins in an otherwise inhospitable
organic solvent, thus enabling these organic phases to be used as pro-
tein extractants. Factors affecting the solubilization effectiveness of
the solvents include charge effects, such as the net charge determined
by the pH relative to the protein isoelectric point; charge distribution
and asymmetry on the protein surface; and the type (anionic or
cationic) of the surfactant used in the reversed micellar phase. Ionic
strength and salt type affect the electrostatic interactions between the
proteins and the surfactants, and affect the sizes of the reversed
micelles. Attachment of affinity ligands to the surfactants has been
demonstrated to lead to enhancements in extraction efficiency and
selectivity [Kelley et al., Biotech. Bioeng., 42, 1199–1208 (1993)].
Product recovery from reversed micellar solutions can often be
attained by simple backextraction, by contacting with an aqueous solu-
tion having salt concentration and pH that disfavors protein solubi-
lization, but this is not always a reliable method. Addition of
cosolvents such as ethyl acetate or alcohols can lead to a disruption of
the micelles and expulsion of the protein species, but this may also
lead to protein denaturation. These additives must be removed by dis-
tillation, e.g., to enable reconstitution of the micellar phase. Temper-
ature increases can similarly lead to product release as a concentrated
aqueous solution. Removal of the water from the reversed micelles by
molecular sieves or silica gel has also been found to cause a precipita-
tion of the protein from the organic phase.
Extraction using liquid emulsion membranes involves the use of a
surface-active agent such as a surfactant to form dispersed droplets
that encapsulate biomolecules. Its economic viability for large-scale
applications is still weak [Patnaik, in Subramanian (ed.), op. cit., vol. 1,
pp. 267–303]. Another less-known method involving the use of a sur-
factant is the foam fractionation method that has seen limited applica-
tions. Ionic fluids have found commercial applications in chemical
reactions by replacing volatile solvents. They are emerging as an envi-
ronmentally friendly solvent replacement in liquid-liquid phase parti-
tioning. Room-temperature ionic liquids are low-melting-point salts
that stay as liquids at room temperature. Partition behavior in a system
involving a room-temperature ionic fluid and an aqueous phase is
influenced by the type of ionic liquid used as well as pH change
(Visser et al., Green Chem., Feb. 1–4, 2000). An ionic fluid was also
reportedly used in the mobile phase for liquid chromatography
[He et al., J. Chromatogr. A, 1007, 39–45 (2003)]. More research
needs to be done in this area to develop this new green technology
[Freemantle, C&EN, 83, 33–38 (2005).]
Aqueous-detergent solutions of appropriate concentration and
temperature can phase-separate to form two phases, one rich in deter-
gents, possibly in the form of micelles, and the other depleted of the
detergent [Pryde and Phillips, Biochem. J., 233, 525–533 (1986)].
SELECTION OF BIOCHEMICAL SEPARATION PROCESSES 20-77
FIG. 20-88 Process scheme for protein extraction in aqueous two-phase systems for the downstream processing of intra-
cellular proteins, incorporating PEG and salt recycling. [Reprinted from Kelly and Hatton in Stephanopoulos (ed.), op. cit.;
adapted from Greve and Kula, op. cit.]
Proteins distribute between the two phases, hydrophobic (e.g., mem-
brane) proteins reporting to the detergent-rich phase and hydrophilic
proteins to the detergent-free phase. Indications are that the size-
exclusion properties of these systems can also be exploited for viral
separations. These systems would be handled in the same way as the
aqueous two-phase systems.
On occasion, for extracellular products, cell separation can be com-
bined with an initial volume reduction and purification step by using
liquid-liquid extraction. This is particularly true for low-molecular-
weight products, and it has been used effectively for antibiotic and vit-
amin recovery. Often scroll decanters can be used for this separation.
The solids are generally kept in suspension (which requires that the
solids be denser than heavy phase), while the organic phase, which
must be lighter than water (cells typically sink in water), is removed.
Experience shows that scrolls are good for handling the variability
seen in fermentation feedstock. Podbielniak rotating-drum extraction
units have been used often, but only when solids are not sticky,
gummy, or flocculated, as they can get stuck in perforations of the
concentric drums, but will actually give stages to the extraction in
short-residence time (temperature-sensitive product). The Karr reci-
procating-plate column can handle large volumes of whole-broth
materials efficiently, and it is amenable to ready scale-up from small
laboratory-scale systems to large plant-scale equipment.
Adsorption Adsorption (see also Sec. 16, “Adsorption and Ion
Exchange”) can be used for the removal of pigments and nucleic
acids, e.g., or can be used for direct adsorption of the desired proteins.
Stirred-batch or expanded-bed operations allow for presence of par-
ticulate matter, but fixed beds are not recommended for unclarified
broths owing to fouling problems. These separations can be effected
through charge, hydrophobic, or affinity interactions between the
species and the adsorbent particles, as in the chromatographic steps
outlined below. The adsorption processes described here are different
from those traditionally ascribed to chromatography in that they do
not rely on packed-bed operations.
In continuous affinity recycle extraction (CARE) operations,
the adsorbent beads are added directly to the cell homogenate, and
the mixture is fed to a microfiltration unit. The beads loaded with the
desired solute are retained by the membrane, and the product is
recovered in a second stage by changing the buffer conditions to dis-
Expanded-bed adsorption (EBA) has gained popularity in biopro-
cessing since its commercial introduction in the 1990s because of its
ability to handle a crude feedstock that contains cells or other particu-
lates. EBA eliminates the need for a dedicated clarification step by
combining solid-liquid separation and adsorption into a single-unit
operation [Hjorth et al., in Subramanian (ed.), op. cit., vol. 1, pp.
199–226; Mattiasson et al., in Ahuja (ed.), op. cit., pp. 417–451]. A
typical EBA operation cycle is illustrated in Fig. 20-89. A bed packed
with an adsorption medium (or a resin), usually spherical particles of
different sizes, is expanded by an upward-flow liquid stream from the
bottom. An unclarified feed is introduced after a stable expansion of
the bed is achieved. Particulates pass through the void spaces between
the resin particles, while the soluble product molecules are adsorbed
by the resin and retained in the column. After a washing step, the
resin particles are left to settle in the column to form a fixed bed. The
product molecules are then eluted out with a mobile phase entering
from the top of the column in a way similar to that in conventional fix-
bed chromatography, to achieve a high-resolution separation. The elu-
tion can also be performed in expanded mode if needed. The
regeneration step in the expanded mode flushes away residual partic-
ulates and refreshes the media for the next cycle. The difference
between EBA and conventional fluidized-bed adsorption lies in the
adsorption resin. In conventional fluidized-bed adsorption, the resin
particles are randomly distributed in the column. In EBA, however,
the resin particles are distributed vertically with large ones near the
column bottom and the small ones near the top. There is no backmixing
along the axial direction of the vertically standing column, thus achiev-
ing adsorption similar to that in a fixed-bed column. The resin particles
have to be prepared to possess a suitable size distribution, or alterna-
tively a distribution based on density differences if the particles sizes are
uniform. An EBA column should have a length typically 3 to 4 times of
the settled bed height to allow for bed expansion. An adjustable adapter
at the top is needed to push the resin downward for elution in the fixed-
bed mode. A proper design of the bottom frit is critical. Its holes must
be smaller enough to retain the smallest resin particles but large enough
to allow the particulates to enter and exit the column freely. In practical
applications, plugging of the frit and nonuniform upward flow tend to
be problematic, especially in columns with large diameters. To mitigate
this problem, some EBA columns use an optimized inlet design and a
mechanical stirrer at the bottom.
The advantages of EBA are its ability to adsorb the soluble product
molecules directly from a cell suspension, a cell homogenate, or a
crude biological fluid containing various particulates, thus making the
“whole-broth processing” concept a reality. EBA eliminates the solid-
liquid separation step (such as microfiltration and centrifugation) and
enables a fast, more compact process requiring fewer steps and less
time. It performs solid removal, concentration, and purification, all in
a single-unit operation. By doing so, it can also minimize the risk of
proteolytic breakdown of the product.
Membrane Ultrafiltration Membrane ultrafiltration is often one
of the favored unit operations used for the isolation and concentration
of biomolecules because they can be easily scaled up to process large
feed volumes at low costs. Toward the end of an ultrafiltration opera-
tion, additional water or buffer is added to facilitate the passage of
20-78 ALTERNATIVE SEPARATION PROCESSES
Resin sedimented Resin expanded Feed loading and washing Fixed-bed elution Regeneration
FIG. 20-89 An operation cycle in expanded-bed adsorption.
smaller molecules. This is known as diafiltration. Diafiltration is espe-
cially helpful in the removal of small contaminating species such as
unspent nutrients including salts and metabolites. Salt removal is usu-
ally necessary if the next step is ion-exchange or reverse-phase chroma-
tography. For a protein that is not very large, two ultrafiltration steps
can be used in sequence. In the first one, the protein ends up in the per-
meate, allowing the removal of large contaminating molecules including
pyrogens and also viral particles. In the second ultrafiltration, the pro-
tein stays in the retentate. This removes small contaminating molecules
and concentrates the feed up to the protein’s solubility limit. Membrane
materials, configurations, and design considerations were discussed ear-
lier in this section. Proper membrane materials must be selected to
avoid undesirable binding with proteins. External fouling, pore block-
age, and internal fouling were discussed by Ghosh (Protein Biosepara-
tion Using Ultrafiltration, Imperial College Press, London, 2003).
The final purification steps are responsible for the removal of the last
traces of impurities. The volume reduction in the earlier stages of the
separation train is necessary to ensure that these high-resolution oper-
ations are not overloaded. Generally, chromatography is used in these
final stages. Electrophoresis can also be used, but since it is rarely
found in process-scale operations, it is not addressed here. The final
product preparation may require removal of solvent and drying, or
lyophilization of the product.
Chromatography Liquid chromatography steps are ubiquitous
in the downstream processing. It is the most widely used downstream
processing operation because of its versatility, high selectivity, and
efficiency, in addition to its adequate scale-up potential based on wide
experience in the biochemical processing industries. As familiarity is
gained with other techniques such as liquid-liquid extraction, they will
begin to find greater favor in the early stages of the separation train,
but are unlikely to replace chromatography in the final stages, where
high purities are needed.
Chromatography is typically a fixed-bed adsorption operation, in
which a column filled with chromatographic packing materials is fed
with the mixture of components to be separated. Apart from size-exclu-
sion (also known as gel-permeation or gel-filtration) chromatography, in
the most commonly practiced industrial processes the solutes are
adsorbed strongly to the packing materials until the bed capacity has
been reached. The column may then be washed to remove impurities in
the interstitial regions of the bed prior to elution of the solutes. This lat-
ter step is accomplished by using buffers or solvents which weaken the
binding interaction of the proteins with the packings, permitting their
recovery in the mobile phase. To minimize product loss of a high-value
product, a small load far below the saturation capacity is applied to the
column. A complete baseline separation can then be achieved after elu-
tion. Gradient elution uses varied modifier strength in the mobile phase
to achieve better separations of more chemical species.
Types of Chromatography Practiced Separation of proteins by
using chromatography can exploit a range of different physical and
chemical properties of the proteins and the chromatography adsorp-
tion media [Janson and Ryden, op. cit.; Scopes, op. cit.; Egerer, in
Finn and Prave (eds.), Biotechnology Focus 1, Hanser Publishers,
Munich, 1988, pp. 95–151]. Parameters that must be considered in
the selection of a chromatographic method include composition of the
feed, the chemical structure and stability of the components, the elec-
tric charge at a defined pH value and the isoelectric point of the pro-
teins, the hydrophilicity and hydrophobicity of the components, and
molecular size. The different types of interactions are illustrated
schematically in Fig. 20-90.
Ion-exchange chromatography relies on the coulombic attraction
between the ionized functional groups of proteins and oppositely
charged functional groups on the chromatographic support. It is used
to separate the product from contaminating species having different
charge characteristics under well-defined eluting conditions, and for
concentration of the product, owing to the high-adsorptive capacity
of most ion-exchange resins and the resolution attainable. Elution is
carried out by using a mobile phase with competing ions or varied
pH. Ion-exchange chromatography is used effectively at the front
end of a downstream processing train for early volume reduction and
The differences in sizes and locations of hydrophobic pockets or
patches on proteins can be exploited in hydrophobic interaction chro-
matography (HIC) and reverse-phase chromatography (RPC); dis-
crimination is based on interactions between the exposed hydrophobic
residues and hydrophobic ligands which are distributed evenly
throughout a hydrophilic porous matrix. As such, the binding charac-
teristics complement those of other chromatographic methods, such
as ion-exchange chromatography. In HIC, the hydrophobic interac-
tions are relatively weak, often driven by salts in high concentration,
and depend primarily on the exposed residues on or near the protein
surface; preservation of the native, biologically active state of the
protein is a desirable feature of HIC. HIC’s popularity is on the rise in
recent years because of this feature. Elution can be achieved differen-
tially by decreasing salt concentration or increasing the concentration
of polarity perturbants (e.g., ethylene glycol) in the eluent.
Reverse-phase chromatography relies on significantly stronger
hydrophobic interactions than in HIC, which can result in unfolding
and exposure of the interior hydrophobic residues, i.e., leads to pro-
tein denaturation and irreversible inactivation; as such, RPC depends
on total hydrophobic residue content. Elution is effected by organic
solvents applied under gradient conditions. RPC is the most com-
monly used analytical chromatographic method due to its ability to
separate a vast array of chemicals with high resolutions. Denaturation
of proteins does not influence the analytical outcome unless protein
precipitation in the mobile phase becomes a problem.
HIC typically uses polymer-based resins with phenyl, butyl, or octyl
ligands while RPC uses silica beads with straight-chain alkanes with 4,
8, or 18 carbons. Larger ligands provide stronger interactions. Poly-
meric beads are used in RPC when basic pH is involved because silica
beads are unstable at such pH. In HIC, the mobile phase remains an
aqueous salt solution, while RPC uses solvent in its mobile phase to
regulate binding. Raising the temperature increases the hydrophobic
interactions at the temperatures commonly encountered in biological
HIC is most effective during the early stages of a purification strat-
egy and has the advantage that sample pretreatment such as dialysis or
desalting after salt precipitation is not usually required. It is also find-
ing increased use as the last high-resolution step to replace gel filtra-
tion. It is a group separation method, and generally 50 percent or
more of extraneous impurities are removed. This method is charac-
terized by high adsorption capacity, good selectivity, and satisfactory
yield of active material.
Despite the intrinsically nonspecific nature of ion-exchange and
reversed-phase/hydrophobic interactions, it is often found that chro-
matographic techniques based on these interactions can exhibit
remarkable resolution. This is attributed to the dynamics of multisite
interactions being different for proteins having differing surface dis-
tributions of hydrophobic and/or ionizable groups.
Size-exclusion chromatography’s (SEC’s) separation mechanism is
based on the sizes and shapes of proteins and impurities. The effective
size of a protein is determined by its steric geometry and solvation
characteristics. Smaller proteins are able to penetrate the small pores
in the beads while large proteins are excluded, making the latter elute
out of column more quickly. To suppress the ion-exchange side
effects, a salt is typically added to the mobile phase. Ammonium car-
bonate or bicarbonate is used if the salt is to be removed by sublima-
tion alone during lyophilization. In rare cases, a solvent at a low
concentration is added to the mobile phase to suppress hydrophobic
interactions between the protein molecules and the stationary phase.
In industrial production processes, SEC columns are used to separate
small molecules from proteins. It is also a choice for desalting and
buffer exchange in the product polishing stage. SEC columns are typ-
ically very large because the feed loads to SEC columns are limited to
3 to 5 percent of the bed volumes. This loading capacity is far less than
those with packings that have binding interactions.
Protein affinity chromatography can be used for the separation of
an individual compound, or a group of structurally similar compounds
from crude-reaction mixtures, fermentation broths, or cell lysates by
exploiting very specific and well-defined molecular interactions
SELECTION OF BIOCHEMICAL SEPARATION PROCESSES 20-79
between the protein and affinity groups immobilized on the packing-
support material. Examples of affinity interactions include antibody-
antigen, hormone-receptor, enzyme-substrate/analog/inhibitor, metal
ion–ligand, and dye-ligand pairs. Monoclonal antibodies are particu-
larly effective as biospecific ligands for the purification of pharmaceu-
tical proteins. Affinity chromatography may be used for the isolation
of a pure product directly from crude fermentation mixtures in a sin-
gle chromatographic step. Immunosorbents should not be subjected
to crude extracts, however, as they are particularly susceptible to foul-
ing and inactivation. Despite its high resolution and the ability to treat
a very dilute feed, affinity chromatography is still costly on the process
scale if protein ligands such as protein A or protein G is to be used, or
a custom affinity matrix is required. Considerable research efforts are
devoted to its development in part due to the increased number of
protein pharmaceuticals produced at low concentrations.
After more than two decades of development, membrane chroma-
tography has emerged as an attractive alternative to packed column
chromatography. Using a porous membrane as the stationary phase in
liquid chromatography has several potential advantages that include a
very high flow rate through a very short and wide bed with only a mod-
est transmembrane pressure drop. Elimination or minimization of dif-
fusional mass-transfer resistance shifts the rate-controlling step to
faster-binding kinetics, resulting in adsorptive separation of proteins
in a fraction of the time required by conventional packed columns. To
achieve sufficient adsorptive separation, it is necessary to use a
medium that binds strongly with target molecules when a very short
flow path is involved. Thus, membrane chromatography typically uses
an affinity membrane, and the combination of membrane chromatog-
raphy with affinity interaction provides high selectivity and fast pro-
cessing for the purification of proteins from dilute feeds. To a much
less extent, ion-exchange, hydrophobic interaction, and reverse-phase
membrane chromatography have also been reported [Charcosset, J.
Chem. Technol. Biotechnol. 71, 95–110 (1998)].
Figure 20-91 shows the interaction between proteins in the
mobile phase and the affinity membrane matrix. The mechanical
strength, hydrophobicity, and ligand density of the membranes can
be engineered through chemical modifications to make them suit-
able for affinity membrane chromatography. A preferred membrane
medium to be prepared for affinity membrane chromatography
should provide (1) desirable physical characteristics such as pore
structure and mechanical strength, allowing fast liquid flow at a
small pressure loss; (2) reactive groups (such as ᎏOH, ᎏNH2,
ᎏSH, ᎏCOOH) for coupling ligands or spacer arms; (3) physical
and chemical stabilities to endure heat or chemical sterilization; and
(4) a nondenaturing matrix to retain protein bioactivity [Zou et al., J.
Biochem. Biophys. Methods, 49, 199–240 (2001)]. Base membranes
20-80 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-90 Schematic illustration of the chromatographic methods most commonly used in downstream processing of
can be chosen from commercial membrane materials including
organic, inorganic, polymeric, and composite materials. The
selected membrane is first activated to create functional groups for
chemical attachment of affinity ligands. If there is steric hindrance
to binding between the immobilized ligand and the target molecule,
a suitable spacer arm is used to bridge the activated membrane sur-
face and the ligand. The ligand should retain its reversible binding
capacity after immobilization onto the support membrane. Affinity
ligands typically fall into two categories: (1) those derived from
enzyme/substrate, antibody/antigen pairs that are capable of very
strong and highly biospecific binding and (2) protein A and protein
G, coenzyme, lectin, dyes, and metal chelates, etc., each capable of
binding with a whole class of molecules. A highly specific ligand pro-
vides an unsurpassed resolution and an ability to handle a large vol-
ume of a dilute feed. However, they are typically fragile and
expensive and may be unavailable off the shelf due to their narrow
applications involving just one or a few molecules that can bind. Elu-
tion can also prove to be a difficult task because some biospecific
bindings can be extremely tenacious. In contrast, a somewhat less
specific ligand has a much wider market and thus is considerably less
expensive. Various membrane cartridges have been used for affinity
membrane chromatography including those with multiple layers of
flat sheet membranes, hollow fibers, and spiral-wound and Chro-
marod membranes (Zou et al., loc. cit.).
Immobilized metal-ion affinity chromatography (IMAC) relies on the
interaction of certain amino acid residues, particularly histidine, cysteine,
and tryptophan, on the surface of the protein with metal ions fixed to the
support by chelation with appropriate chelating compounds, invariably
derivatives of iminodiacetic acid. Commonly used metal ions are Cu2+
. Despite its relative complexity in terms of the num-
ber of factors that influence the process, IMAC is beginning to find
industrial applications. The choice of chelating group, metal ion, pH, and
buffer constituents will determine the adsorption and desorption charac-
teristics. Elution can be effected by several methods, including pH gra-
dient, competitive ligands, organic solvents, and chelating agents.
Following removal of unbound materials in the column by washing,
the bound substances are recovered by changing conditions to favor
desorption. A gradient or stepwise reduction in pH is often suitable.
Otherwise, one can use competitive elution with a gradient of increas-
ing concentration. IMAC eluting agents include ammonium chloride,
glycine, histamine, histidine, or imidazol. Inclusion of a chelating
agent such as EDTA in the eluent will allow all proteins to be eluted
indiscriminately along with the metal ion.
Chromatographic Development The basic concepts of chro-
matographic separations are described elsewhere in this handbook.
Proteins differ from small solutes in that the large number of charged
and/or hydrophobic residues on the protein surface provide multiple
binding sites, which ensure stronger binding of the proteins to the
adsorbents, as well as some discrimination based on the surface distri-
bution of amino acid residues. The proteins are recovered by elution
with a buffer that reduces the strength of this binding and permits the
proteins to be swept out of the column with the buffer solution. In iso-
cratic elution, the buffer concentration is kept constant during the
elution period. Since the different proteins may have significantly dif-
ferent adsorption isotherms, the recovery may not be complete, or it
may take excessive processing time and cause excessive band spread-
ing to recover all proteins from the column. In gradient elution oper-
ations, the composition of the mobile phase is changed during the
process to decrease the binding strength of the proteins successively,
the more loosely bound proteins being removed first before the elu-
ent is strengthened to enable recovery of the more strongly adsorbed
species. The change in eluent composition can be gradual and contin-
uous, or it can be stepwise. Industrially, in large-scale columns it is dif-
ficult to maintain a continuous gradient owing to difficulties in fluid
distribution, and thus stepwise changes are still used. In some adsorp-
tion modes, the protein can be recovered by the successive addition of
competing compounds to displace the adsorbed proteins. In all cases,
the product is eluted as a chromatographic peak, with some possible
overlap between adjacent product peaks.
Displacement chromatography relies on a different mode of elution.
Here a displacer that is more strongly adsorbed than any of the pro-
teins is introduced with the mobile phase. As the displacer concentra-
tion front develops, it pushes the proteins ahead of itself. The more
strongly adsorbed proteins then act as displacers for the less strongly
bound proteins, and so on. This leads to the development of a displacer
train in which the different molecules are eluted from the column in
abutting roughly rectangular peaks in the reverse order of their bind-
ing strength with the column’s stationary phase. Many displacers have
been developed in the past two decades including both high- and low-
molecular-weight molecules. Displacement chromatography may be
practical as an earlier chromatography step in a downstream process
when baseline separations are not necessary. It has been considered for
some industrial processes. However, displacer reuse and possible con-
tamination of the protein product in addition to its inherent inability to
achieve a clear-cut baseline separation make the use of displacement
chromatography still a challenge. So far, no FDA-approved process
exists [Shukla and Cramer, in Ahuja (ed.), op. cit., pp. 379–415].
For efficient adsorption it is advisable to equilibrate both the col-
umn and the sample with the optimum buffer for binding. Prior to
this, the column must be cleaned to remove tightly bound impurities
by increasing the salt concentration beyond that used in the product
elution stages. At the finish of cleaning operation, the column should
be washed with several column volumes of the starting buffer to
remove remaining adsorbed material. In desorption, it is necessary to
drive the favored binding equilibrium for the adsorbed substance
from the stationary to the mobile phase. Ligand-protein interactions
are generally a combination of electrostatic, hydrophobic, and hydro-
gen bonds, and the relative importance of each of these and the
degree of stability of the bound protein must be considered in select-
ing appropriate elution conditions; frequently compromises must be
made. Gradient elution often gives good results.
Changes in pH or ionic strength are generally nonspecific in elution
performance; ionic strength increases are effective when the protein
binding is predominantly electrostatic, as in IEC. Polarity changes are
effective when hydrophobic interactions play the primary role in pro-
tein binding. By reducing the polarity of the eluting mobile phase, this
phase becomes a more thermodynamically favorable environment for
the protein than adsorption to the packing support. A chaotropic salt
(KSCN, KCNO, KI in range of 1 to 3 M) or denaturing agent (urea,
guanidine HCl; 3 to 4 M) in the buffer can also lead to enhanced des-
orption. For the most hydrophobic proteins (e.g., membrane pro-
teins) one can use detergents just below their critical micelle
concentrations to solubilize the proteins and strip them from the
packing surface. Specific elution requires more selective eluents. Pro-
teins can be desorbed from ligands by competitive binding of the
eluting agent (low concentration of 5 to 100 mM) either to the ligand
or to the protein.
Specific eluents are most frequently used with group-specific
adsorbents since selectivity is greatly increased in the elution step.
SELECTION OF BIOCHEMICAL SEPARATION PROCESSES 20-81
FIG. 20-91 Illustration of the interactions between proteins and membrane
matrix in affinity membrane chromatography.
The effectiveness of the elution step can be tailored by using a single
eluent, pulses of different eluents, or eluent gradients. These systems
are generally characterized by mild desorption conditions. If the elut-
ing agent is bound to the protein, it can be dissociated by desalting on
a gel filtration column or by diafiltration.
Column Packings The quality of the separation obtained in
chromatographic separations will depend on the capacity, selectivity,
and hydraulic properties of the stationary phase, which usually con-
sists of porous beads of hydrophilic polymers filled with the solvent.
The xerogels (e.g., crosslinked dextran) shrink and swell depending on
solvent conditions, while aerogels have sizes independent of solution
conditions. A range of materials are used for the manufacture of gel
beads, classified according to whether they are inorganic, synthetic, or
polysaccharides. The most widely used materials are based on neutral
polysaccharides and polyacrylamide. Cellulose gels, such as crosslinked
dextran, are generally used as gel filtration media, but can also be used
as a matrix for ion exchangers. The primary use of these gels is for
desalting and buffer exchange of protein solutions, as nowadays frac-
tionation by gel filtration is performed largely with composite gel
matrices. Agarose, a low-charge fraction of the seaweed polysaccha-
ride agar, is a widely used packing material.
Microporous gels made by point crosslinking dextran or polyacry-
lamides are used for molecular-sieve separations such as size-exclusion
chromatography and gel filtration, but are generally too soft at the porosi-
ties required for efficient protein chromatography. Macroporous gels are
most often obtained from aggregated and physically crosslinked poly-
mers. Examples include agarose, macroreticular polyacrylamide, silica,
and synthetic polymers. These gels are good for ion-exchange and affinity
chromatography as well as for other adsorption chromatography tech-
niques. Composite gels, in which the microporous gel is introduced into
the pores of macroreticular gels, combine the advantages of both types.
High matrix rigidity is offered by porous inorganic silica, which can
be derivatized to enhance its compatibility with proteins, but it is
unstable at alkaline pH. Hydroxyapatite particles have high selectivity
for a wide range of proteins and nucleic acids.
Traditional porous media have pore sizes typically in the range of
100 to 300 Å. To reduce intraparticle diffusion for fast chromatog-
raphy, column packing materials can be made either nonporous or
extremely porous. Gustavsson and Larsson [in Hatti-Kaul and Matti-
asson (eds), op. cit., pp. 423–454] discussed various chromatography
media for fast chromatography of proteins. Nonporous particles such
as the popular 5-µm modified silica beads for reverse-phase chroma-
tography are excellent for analytical applications. However, due to
their limited binding sites that exist only on the outer surface of the
particles, nonporous particles are not suitable for preparative- or
large-scale applications. Gigaporous media are gaining momentum in
recent years. They are particles with pore sizes above 1000 Å. Some
have large enough interconnecting pores in the range of 4000 to 8000
Å that allow even convective flow inside the particles. POROS®
fusion chromatography media are the first commercial products in
this category introduced in the 1980s. A few more products are being
media are synthesized in two steps. Nano-
size subparticles are first synthesized and then polymerized to form
10- to 50-µm particles in a second step. In recent years, advances in
suspension polymerization produced a new type of integral giga-
porous media with improved physical strength. The spherical particles
are formed in a single polymerization step. An oil phase (dispersed
phase) consisting of a monomer (such as styrene), a crosslinking agent
(such as divinylbenzene), an initiator, a diluent, and a special porogen
is used. By dispersing the oil phase in a water phase containing a sta-
bilizer, a suspension can be obtained. The suspension polymerization
is carried out at an elevated temperature above the decomposition
temperature of the initiator to obtain the polymer particles. The dilu-
ent and porogen in the particles form smaller diffusion pores and
much larger through-pores, respectively. Figure 20-92 shows such a
particle with both conventional diffusion pores and gigaporous
through-pores. The through-pores allow convective flow to improve
mass transfer, and the diffusion pores provide an overall large surface
area required for a large binding capacity.
In the scale-up of a chromatographic column with conventional
packing media, the column length scale-up is limited because of the
subsequently increased column pressure. To increase the feed load,
the column is typically scaled up by enlarging the column diameter
after the column length reaches a certain limit. The resultant
pancake-shaped column leads to a deteriorating resolution due to
poor flow distribution in the column cross-section. With rigid giga-
porous particles, axial direction scale-up becomes possible because
the flow rate can be up to 40 times higher than that for the conven-
tional media. This kind of scale-up is far superior because of the
enhanced resolution together with increased feed capacity when the
column axial length is increased. With the advances made in rigid
gigaporous media, chromatography columns can potentially process
large volumes of very dilute feeds when combined with strong binding
kinetics such as those involved in affinity chromatography. This may
have a significant impact on the overall downstream process design.
Alternative Chromatographic Columns Commercial radial-
flow chromatography (RFC) columns first appeared in the mid-1980s.
RFC is an alternative to the conventional axial-flow chromatography
for preparative- and large-scale applications. In a RFC column, the
mobile phase flows in the radial direction rather than the axial direc-
tion. Computer simulation proves that RFC is somewhat equivalent to
a pancakelike axial-flow column [Gu, in Flickinger and Drew (eds.),
20-82 ALTERNATIVE SEPARATION PROCESSES
FIG. 20-92 SEM image of a poly(styrene-co-divinylbenzene) gigaporous particle synthesized from suspension poly-
merization and schematic of a gigaporous particle showing through-pores and diffusion pores [Gu et al., China Partic-
uology, 3, 349 (2005)].
op. cit., pp. 627–639]. Both configurations offer a short flow path and
thus a low pressure drop.
Continuous chromatography can be achieved by using a simulated
moving-bed (SMB) process shown in Sec. 16, in which several
columns are linked with a switching device to simulate a continuous
countercurrent flow process. This design maximizes productivity
while minimizing eluent consumption [Nicoud, in Ahuja (ed.), op. cit.,
pp. 475–509]. It is suitable for a simple feed that results in only a lim-
ited fraction [Imamoglu, in Freitag (ed.), op. cit., pp. 212–231]. By
switching the feed, eluant input positions and raffinate, extract output
positions periodically in a series of columns to simulated countercur-
rent movement of the liquid phase and the solid (resin) phase, contin-
uous chromatography is obtained through the use of multiple columns
in series. A four-zone SMB consisting of four columns is capable of
producing two fractions as a pair of raffinate and exact. To have two
pairs (with a total of four fractions) of raffinate and extract outputs an
eight-zone SMB system is needed. Another alternative design
described in Sec. 16 is the so-called annual-flow column that rotates
continuously in the angular direction. Despite its obvious advantage of
being straightforwardly continuous, it suffers from angular dispersion
and reduced bed volume. This design so far has seen very limited
application since its commercial introduction in 1999 [Wolfgang and
Prior, in Freitag (ed.), op. cit., pp. 233–255].
In the past decade, monolithic columns have gained popularity for
analytical applications. Instead of using discrete packing particles, a
whole polymer block is used as a column. Their continuous homoge-
neous structure provides fast mass-transfer rates and very high flow
rates inside the column [Strancar et al., in Freitag (ed.), op. cit., pp.
50–85]. Thin monolithic disks with affinity binding are capable of fast
chromatographic separations. They act much as affinity membrane
chromatography cartridges do. To utilize long monolithic columns for
process-scale separations, a breakthrough in column fabrication is
needed to produce large columns suitable for commercial applications.
Sequencing of Chromatography Steps The sequence of
chromatographic steps used in a protein purification train should be
designed such that the more robust techniques are used first, to
obtain some volume reduction (concentration effect) and to remove
major impurities that might foul subsequent units; these robust units
should have high chemical and physical resistance to enable efficient
regeneration and cleaning, and they should be of low material cost.
These steps should be followed by the more sensitive and selective
operations, sequenced such that buffer changes and concentration
steps between applications to chromatographic columns are
avoided. Frequently, ion-exchange chromatography is used as the
first step. The elution peaks from such columns can be applied
directly to hydrophobic interaction chromatographic columns or to a
gel filtration unit, without the need for desalting of the solution
between applications. These columns can also be used as desalting
operations, and the buffers used to elute the columns can be
selected to permit direct application of the eluted peaks to the next
Factors to be considered in making the selection of chromatog-
raphy processing steps are cost, sample volume, protein concentration
and sample viscosity, degree of purity of protein product, presence of
nucleic acids, pyrogens, and proteolytic enzymes. Ease with which
different types of adsorbents can be washed free from adsorbed cont-
aminants and denatured proteins must also be considered.
Scale-up of Liquid Chromatography The chromatography
columns in downstream processing typically are operated in the non-
linear region due to a concentrated or overloaded feed. Their scale-up
remains a challenging task. There are two general approaches: (1) the
rule-based method using equations for column resizing (Ladisch, op.
cit., pp. 299–448) and (2) the computer simulation method using rate
models (Gu, Mathematical Modeling and Scale-Up of Liquid Chro-
matography, Springer-Verlag, Berlin-New York, 1995) with simula-
tion software such as Chromulator®
to predict column performance
for a particular column setting and operating conditions.
PRODUCT POLISHING AND FORMULATION
The product from the final purification unit operation is typically in a
liquid fraction containing water, a solvent, or a buffer. Based on the
requirement for the final product, they may need to be removed. A
solid protein is usually far more stable with a much longer shelf life.
Product formulation may also require an excipient to be added. Thus,
additional unit operations are needed after the final purification step.
Lyophilization and Drying After the last high-performance
purification steps it is usually necessary to prepare the finished product
for special applications. For instance, final enzyme products are often
required in the form of a dry powder to provide for stability and ease of
handling, while pharmaceutical preparations also require high purity,
stability during formulation, absence of microbial load, and extended
shelf life. This product formulation step may involve drying of the final
products by freeze drying, spray drying, fluidized-bed drying, or crys-
tallization (Golker, in Stephanopoulos, op. cit., pp. 695–714). Crystal-
lization can also serve as an economical purification step [Lee and Kim,
in Hatti-Kaul and Mattiasson (ed.), op. cit., pp. 277–320] in addition to
its role as a unit operation for product polishing.
Freeze drying, or lyophilization, is normally reserved for temperature-
sensitive materials such as vaccines, enzymes, microorganisms, and
therapeutic proteins, as it can account for a significant portion of total
production cost. This process is characterized by three distinct steps,
beginning with freezing of the product solution, followed by water
removal by sublimation in a primary drying step, and ending with sec-
ondary drying by heating to remove residual moisture.
Freezing is carried out on cooled plates in trays or with the product
distributed as small particles on a drum cooler; by dropping the prod-
uct solution in liquid nitrogen or some other cooling liquid; by
cospraying with liquid CO2 or liquid nitrogen; or by freezing with cir-
culating cold air. The properties of the freeze-dried product, such as
texture and ease of rehydration, depend on the size and shape of the
ice crystals formed, which in turn depend on the degree of under-
cooling. It is customary to cool below the lowest equilibrium eutectic
temperature of the mixture, although many multicomponent mixtures
do not exhibit eutectic points. Freezing should be rapid to avoid
effects from local concentration gradients. Removal of water from
solution by the formation of ice crystals leads to changes in salt con-
centration and pH, as well as enhanced concentration of the product,
in the remaining solution; this in turn can enhance reaction rates, and
even reaction order can change, resulting in cold denaturation of the
product. If the feed contains a solvent and an acid, the solvent tends
to sublimate faster than the acid, causing acidic damage to the protein.
With a high initial protein concentration the freeze concentration fac-
tor and the amount of ice formed will be reduced, resulting in greater
product stability. For aseptic processing, direct freezing in the freeze-
drying plant ensures easier loading of the solution after filtration than
if it is transferred separately from remote freezers.
In the primary drying step, heat of sublimation is supplied by con-
tact, conduction, or radiation to the sublimation front. It is important
to avoid partial melting of the ice layer. Many pharmaceutical prepa-
rations dried in ampoules are placed on heated shelves. The drying
time depends on the quality of ice crystals, indicating the importance
of controlling the freezing process; smaller crystals offer higher inter-
facial areas for heat and mass transfer, but larger crystals provide
pores for diffusion of vapor away from the sublimation front.
A high percentage of water remains after the sublimation process,
present as adsorbed water, water of hydration, or dissolved in the dry
amorphous solid; this is difficult to remove. Usually, shelf temperature
is increased to 25 to 40°C and chamber pressure is lowered as far as
possible. This still does not result in complete drying, however, which
can be achieved only by using even higher temperatures, at which
point thermally induced product degradation can occur.
Excipients can be used to improve stability and prevent deterioration
and inactivation of biomolecules through structural changes such as dis-
sociation from multimeric states into subunits, decrease in α-helical
content accompanied by an increase in β-sheet structure, or complete
unfolding of helical structure. These are added prior to the freeze-dry-
ing process. Examples of these protective agents include sugars, sugar
derivatives, and various amino acids, as well as polymers such as dex-
tran, polyvinyl pyrrolidone, hydroxyethyl starch, and polyethylene gly-
col. Some excipients, the lyoprotectants, provide protection during
freezing, drying, and storage, while others, the cryoprotectants, offer
protection only during the freezing process. Spray drying can use up to
50 percent less energy than freeze-drying operations and finds applica-
SELECTION OF BIOCHEMICAL SEPARATION PROCESSES 20-83
tion in the production of enzymes used as industrial catalysts, as addi-
tives for washing detergents, and as the last step in the production of
single-cell protein. The product is usually fed to the dryer as a solution,
a suspension, or a free-flowing wet substance.
Spray drying is an adiabatic process, the energy being provided by
hot gas (usually hot air) at temperatures between 120 and 400°C.
Product stability is ensured by a very short drying time in the spray-
drying equipment, typically in the subsecond to second range, which
limits exposure to the elevated temperatures in the dryer. Protection
can be offered by addition of additives (e.g., galactomannan, polyvinyl
pyrrolidone, methyl cellulose, cellulose). The spray-drying process
requires dispersion of the feed as small droplets to provide a large
heat- and mass-transfer area. The dispersion of liquid is attained by
using rotating disks, different types of nozzles, or ultrasound, and is
affected by interfacial tension, density, and dynamic viscosity of the
feed solution, as well as the temperature and relative velocities of the
liquid and air in the mixing zone. Rotating-disk atomizers operate at
4000 to 50,000 rpm to generate the centrifugal forces needed for dis-
persion of the liquid phase; typical droplet sizes of 25 to 950 µm are
obtained. These atomizers are especially suitable for dispersing sus-
pensions that would tend to clog nozzles. For processing under asep-
tic conditions, the spray drier must be connected to a filling line that
allows aseptic handling of the product.
Thickeners and binders such as acacia, agar, starch, sodium alig-
nate, gelatin, methyl cellulose, bentonite, and silica are used to
improve product stability and enhance the convenience of the admin-
istration of a liquid formulation. Surface-active agents, colors, flavors;
and preservatives may also be used in the final formulation (Garcia
et al., Bioseparation Process Science, Blackwell Science, Malden,
Mass., 1999, p. 374).
INTEGRATION OF UNIT OPERATIONS IN
Generally speaking, a typical downstream process consisting of four
stages: removal of isolubles, isolation, purification, and polishing (Bel-
ter et al., op. cit., p. 5). Cell disruption is required for intracellular
products. One or two, and sometimes more chromatography steps will
serve as the center-stage unit operations. The steps before them serve
the purposes of feed volume reduction and removal of the majority of
impurities. The steps after them are polishing and formulation opera-
tions. Based on these general outlines, a few rules of thumb may be
used (Harrison et al., op. cit., p. 322; Garcia et al., op. cit., p. 358): (1)
Reduce the feed volume early in the process, (2) remove the most
abundant impurities or the easy-to-remove impurities first, (3) reduce
the amount of impurities as much as possible before the delicate high-
resolution chromatography steps, and (4) sequence the unit opera-
tions that exploit different separation mechanisms.
The purification of proteins to be used for therapeutic purposes
presents more than just the technical problems associated with the
separation process. Owing to the complex nature and intricate three-
dimensional structure, the routine determination of protein structure
as a quality control tool, particularly in its final medium for use, is not
well established. In addition, the complex nature of the human
immune system allows for even minor quantities of impurities and
contaminants to be biologically active. Thus, regulation of biologics
production has resulted in the concept of the process defining the
product since even small and inadvertent changes in the process may
affect the safety and efficacy of the product. Indeed, it is generally
acknowledged that even trace amounts of contaminants introduced
from other processes, or contaminants resulting from improper
equipment cleaning, can compromise the product. From a regulatory
perspective, then, operations should be chosen for more than just effi-
ciency. The consistency of the unit operation, particularly in the face
of potentially variable feed from the culture/fermentation process, is
the cornerstone of the process definition. Operations that lack robust-
ness or are subject to significant variation should not be considered.
Another aspect of process definition is the ability to quantify the oper-
ation’s performance. Finally, the ease with which the equipment can
be cleaned in a verifiable manner should play a role in unit operation
selection. Obviously, certain unit operations are favored over others
because they are easier to validate. Process validation was covered in
a book edited by Subramanian (op. cit., vol. 2, pp. 379–460).
Keep in mind that some unit operations are not as scalable as
others. The evolution of a bench-scale process to production scale will
see changes in the types and number of unit operations selected. To
configure an effective and efficient bioseparations process, a thorough
understanding of the various unit operations in downstream process-
ing described above is a prerequisite. Different process-scale and
purity requirements can necessitate changes and may result in quite
different configurations for the same product. Existing process exam-
ples and past experiences help greatly. Due to regulatory restrictions,
once a process is approved for a biopharmaceutical, any change in unit
operations requires time-consuming and costly new regulatory
approval. This means that an optimized process design is much
desired before seeking approval. The involvement of biochemical
engineers early in the design process is highly recommended.
INTEGRATION OF UPSTREAM
AND DOWNSTREAM OPERATIONS
Upstream fermentation (or cell culture) has a direct impact on the
design and optimization of its downstream process. Different media or
different operating conditions in fermentation result in a different feed
for the downstream process. In the development of new products,
optimization of the fermentation medium for titer only often ignores
the consequences of the medium properties on subsequent down-
stream processing steps such as filtration and chromatography. It is
imperative, therefore, that there be effective communication and
understanding between workers on the upstream and downstream
phases of the product development if rational tradeoffs are to be made
to ensure overall optimality of the process. One example is to make
the conscious decision, in collaboration with those responsible for the
downstream operations, of whether to produce a protein in an
unfolded form or in its native folded form. The purification of the
aggregated unfolded proteins is simpler than that of the native pro-
tein, but the refolding process itself to obtain the product in its final
form may lack scalability or certainty in its method development.
In some instances, careful consideration of the conditions used in
the fermentation process, or manipulation of the genetic makeup of
the host, can simplify and even eliminate some unit operations in the
downstream processing sequence [Kelley and Hatton, Bioseparation,
1, 303–349 (1991)]. Some of the advances made in this area are the
engineering of strains of E. coli to allow the inducible expression of
lytic enzymes capable of disrupting the wall from within for the release
of intracellular protein products, the use of secretion vectors for the
expression of proteins in bacterial production systems. Fusion proteins
can be genetically engineered to attach an extra peptide or protein that
can bind with an affinity chromatography medium [Whitmarsh and
Hornby, in Street (ed.), op. cit., pp. 163–177]. This can enhance the
purification of an otherwise difficult to purify protein greatly. The cell
culture medium can be selected to avoid components that can hinder
subsequent purification procedures. Integration of the fermentation
and initial separation/purification steps in a single operation can also
lead to enhanced productivity, particularly when the product can be
removed as it is formed to prevent its proteolytic destruction by the
proteases which are frequently the by-product of fermentation
processes. The introduction of a solvent directly to the fermentation
medium (e.g., phase-forming polymers), the continuous removal of
products by using ultrafiltration membranes, and the use of continuous
fluidized-bed operations are examples of this integration.
Process economics for biological products was discussed by Harri-
son et al. (op. cit., pp. 334–369) and Datar and Rosen [in Asenjo (ed.),
Separation Processes in Biotechnology, Dekker, New York, 1990, pp.
741–793] at length, and also by Ladisch (op. cit., pp. 401–430). They
provided some illustrative examples with cost analyses. Bioprocess
design software can also prove helpful in the overall design process
(Harrison et al., op. cit., pp. 343–369).
20-84 ALTERNATIVE SEPARATION PROCESSES