Chapter 7
Innovative separation methods in
bioprocessing
J. A. ASENJO, Biochemical Engineering Laboratory, Department of Food Science and
Technology, The University of Reading, Reading RG6 6AP, UK and J. B.
CHAUDHURI, School of Chemical Engineering, University of Bath, Bath BA2 7AY,
UK
7.1 INTRODUCTION
Discoveries and achievements in modern biology and recombinant DNA technology in
the last few years have resulted in the development of a number of new therapeutics for
human use such as insulin, human growth hormone (hGH), tissue plasminogen activator
(tPA) for cardiac disease, erythropoietin (EPO) and hepatitis B vaccine and thus the
possibility of their industrial large-scale production. This poses a tremendous challenge
for the chemical and biochemical engineer in terms of developing efficient separation
processes for these new proteins. As they are intended for human use the levels of purity
required are of the order of 99.9% or 99.98% or even higher (depending on dosage) and
they have to be separated from a very large number of contaminants, other proteins,
nucleic acids, polysaccharides and many other components present in the cell culture or
cell lysate used to manufacture these proteins. Competitive advantage in production
depends not only on innovations in molecular biology and other areas of basic biological
sciences but also on innovation and optimisation of separation and downstream proc-
esses.
The main issues important for the development of novel separation techniques to give
improved resolution, simplicity, speed, ease of scale-up and possibly continuous
operation are presented and discussed. The assessment of the state of the art as well as
promising future developments concentrate on the separation and purification of proteins
from complex mixtures. The present trend to develop techniques that exploit fundamental
physicochemical principles more efficiently is emphasised. This includes the analysis of
the physicochemical properties of proteins such as PI, charge as a function of pH,
biological affinity (including metal ion and dye affinity), hydrophobicity and size and its
180
relation to efficiency in a bioseparation. Some properties (e.g. charge and affinity) can
show extremely high resolution in purification operations, whereas others (e.g. molecular
weight) show much lower resolution.
7.2 SYSTEM CHARACTERISTICS
7.2.1 Physicochemical basis for separation operations
Development of new and efficient separation processes will be based on more effectively
exploiting differences in the actual physicochemical properties of the product such as
surface charge/titration curve, surface hydrophobicity, molecular weight, biospecificity
towards certain ligands (e.g. metal ions, dyes), PI and stability, compared to those of the
contaminant components in the crude broth. The main physicochemical factors involved
in the development of separation processes are shown in Table 7.1 (Asenjo, 1993).
J. A. Asenjo and J. B. Chaudhuri
Table 7.1. Physicochemical basis for the development of separation
processes
Physicochemical basis Separation process
Charge Ion-exchange chromatography
Electrodialysis
Aqueous two-phase partitioning
Reverse micelle extraction
Hydrophobicity Hydrophobic interaction chromatography
Reversed phase chromatography
Precipitation
Aqueous two-phase partitioning
Specific binding Affinity chromatography
Size Gel filtration
Ultrafiltration
Dialysis
Electric mobility Electrophoresis
Isoelectric point Chromatofocusing
Isoelectric focusing
Sedimentation rate Centrifugation
Surface activity Adsorption
Solubility Solid-liquid extraction
Foam fractionation
Supercritical fluid extraction
(From Asenjo, 1993)
Innovative separation methods in bioprocessing 18 1
7.2.2 Kinetics and mass transfer
The physical behaviour of the system has an effect on the development of novel
separation processes. Processes can be divided into equilibrium and rate processes. In
equilibrium processes selective separation depends on the attainment of a favourable
equilibrium state. This, for example, includes liquid-liquid extraction and ion exchange
chromatography, Rate processes, on the other hand, separate different proteins on the
basis of their response to an imposed field (such as an electric field). Mobility and other
similar properties determine the selectivity of this type of operation; a successful process
is one in which the proteins have markedly different mobilities (e.g. electrophoresis).
In a number of protein separation processes the residence time in the reactor is
insufficient for equilibrium to be achieved and the kinetics of adsorption play an
important role for example in affinity chromatography ana in the CARE (continuous
adsorption recycle extraction) process. New developments in materials have recently
shown dramatic advances in overcoming mass transfer limitations in processes such as
perfusion and membrane chromatography and adsorption resulting in extremely fast
separations, Some recent examples of novel techniques, which exploit the principles
discussed above and provide useful analyses for optimal design of operations, include
expanded bed (fluidised bed) adsorption of proteins, which allows direct broth extraction;
cross-flow electrofiltration of disrupted microbial cells and for improved ultrafiltration of
proteins; mathematical modelling of partitioning and phase behaviour in liquid-liquid
extraction; mathematical modelling of chromatographic columns; perfusion and
membrane chromatography; and advanced reversed phase chromatography using HPLC.
The potential for scale-up of many of these systems is analysed and discussed.
7.3 LIQUID-LIQUID EXTRACTION: INTRODUCTION
Liquid-liquid extraction as a technology has been used in the antibiotics industry for
several decades and it is now beginning to be recognised as a potentially useful
separation step in protein recovery and separation, particularly because it can readily be
scaled-up and can, if necessary, be operated on a continuous basis. The physicochemical
factors of the protein that determine partitioning are also starting to be understood. It is a
reasonably high-capacity process and can offer good selectivity for the desired protein
product. However, poor solubility of the large protein molecules in typical organic
solvents restricts the range of solvents available for use in such a separation process.
Two classes of solvents that appear to offer advantages for protein recovery for protein
separations are aqueous polymer/salt (in some cases also polymer/polymer) systems and
reverse micellar solutions. In both cases two phases are formed and the separation
exploits the difference in partitioning of the proteins in the feed and extraction phases. In
the aqueous polymer/salt separation systems the partitioning of the protein occurs
between two immiscible aqueous phases; one rich in a polymer (usually polyethylene
glycol, PEG) and the other in a salt (e.g. phosphate or sulphate). These systems show a
non-denaturing solvent environment, small interfacial resistance to mass transfer,
relatively high protein capacity and high selectivity. On the other hand reverse micelles
exploit the solubilising properties of surfactants that can aggregate in organic solvents to
form so-called inverted or reverse micelles. These aggregates consist of a polar core of
182
water and the solubilised protein stabilised by a surfactant shell layer. For protein
extraction, one phase is the aqueous feed solution, the other the reversed micellar phase
that acts as the extractant. They have several of the advantages quoted for aqueous two-
phase systems.
The suitability of using foam separation as well as gas aphrons as novel separation
techniques for proteins are presently under investigation.
7.3.1 Aqueous two-phase separation
Partitioning in two aqueous phases can be used for the separation of proteins from cell
debris as well as for purification from other proteins. Partitioning can be done in a single
step or as a multistage process. Differences in partition coefficients, however, between
the different proteins can be high, hence one step tends to be sufficient (usually one for
extraction and one for elution or back-extraction). The use of affinity partitioning can
greatly enhance the specificity of the extraction. A typical process for extraction of a
protein into the top PEG phase in a first stage and the back extraction into a bottom salt
phase (e.g. phosphate or sulphate) in a second ‘back extraction’ step from a cell
homogenate that includes recycle of the PEG phase is shown in Fig. 7.1 (Hustedt et al.,
1985).
J. A. Asenjo and J. B. Chaudhuri
....................................................................
Fig. 7.1. Scheme of enzyme purification by liquid-liquid extraction. The cells are disrupted by
wet milling, and after passing through a heat exchanger, PEG and salts are added into the process
stream of broken cells. After mixing and obtaining of equilibrium the phase system is separated,
the outflowing bottom phase is going to waste. The product-containing PEG-rich top phase goes
to a second mixer after addition of more salt to the process stream. The product is recovered in
the resulting bottom phase while the concentrated PEG solution (upper phase) goes to waste or is
recycled. (From Hustedt et al., 1985)
Innovative separation methods in bioprocessing 183
Most soluble and particulate material partitions to the lower, more polar (e.g. salt)
phase and the protein of interest partitions to the top less polar phase, usually PEG.
Separation of actual proteins in such systems is based on manipulating the partition
coefficient (K) by altering parameters such as average molecular weight of the polymer,
type of phase forming salt used for the heavy phase, the types of ions included in the
system and ionic strength of added salts (e.g. NaCI) (Schmidt et al., 1994). Figure 7.2
shows that the partition coefficient of a-amylase is a strong function of the presence of
NaCl in a PEG/sulphate system. For extraction of the a-amylase from its contaminants a
high concentration of NaCl is used in the first extraction stage, whereas a low concentra-
tion of NaCl in the back-extraction stage will allow the recovery of a-amylase into the
bottom sulphate phase as shown in Fig. 7.1.
4-
2-
-4 I I I I I
0 2 4 6 a 10
Concentration of NaCl (Yo w/W)
Fig. 7.2. Partition behaviour of a-amylase (log KJ and contaminant protein (log K,) from
industrial supernatant from E. subtilis fermentation in PEG 4000/Sulphate systems as a function
of added NaCl concentration at pH 7 and a phase volume ratio of 1.
The partition coefficient (K) is defined as the concentration of a particular protein in
the lighter phase divided by the concentration in the heavier phase. The main factors that
determine partition depend on the type of system used:
(1)
Hydrophobicity. Differences in the surface hydrophobicity between proteins are
exploited when partitioning them in PEG/salt two-phase systems. Typical systems
that exploit a protein’s hydrophobicity are PEG/phosphate and PEG/sulphate with
addition of a high concentration of NaCl (e.g. 10%).
Size-dependent partition. Molecular size of the proteins or surface area of the
particles to be partitioned is the dominating factor. It has been shown that for
PEG/Dextran systems a protein’s molecular weight is inversely proportional to its
partition coefficient.
Electrochemical. Electrical potential between the phases is used to separate
molecules or particles according to their charge. This is demonstrated in
(2)
(3)
1.0
Y
g 0.5
-
0
-
-
I I
0.4
0.2
0.0
E .
9
v) -0.2
0, 0
-
-0.4
-0.6
-0.8
-
-
I
-
-
-
-
-
I I
-4
-
Innovative separation methods in bioprocessing 187
Organic phase
4V&
.t nv
-
Reverse
4 ' L micelle
?fiC%4k/
h
3 nc
L.UVVaUL. V UdU UU!iiCUY u
Protein
a
0
a
Aqueous phase
Fig. 7.6. Protein partition into reverse micelles.
There are two techniques for transferring proteins into the micellar phase. The most
widely used method involves extraction of the protein with a biphasic liquid system, i.e.
liquid-liquid extraction. One phase is the aqueous solution of the protein, and the other
the organic micellar solution, usually in equal volume. By gently shaking the two phases,
the protein partitions from the aqueous into the micellar phase. In the second method,
solid state extraction of the protein, the protein powder is suspended in the micellar phase
and gently stirred.
The protein solubilised in the reverse micellar solution can be transferred back into an
aqueous solution, by contacting the micellar solution with an aqueous solution containing
a high concentration of a particular salt (KC1, CaC12), which has the capability to ex-
change with the protein in the micelles.
The basic idea is that the process of protein extraction by reverse micelles can be made
specific (Le. tailored to a specific protein) and efficient (Le. high extraction yield) by
controlling the micellar parameters such as the water content, the type and concentration
of surfactant, the type and concentration of salt, and the pH.
Leser et al. (1986) examined the transfer of ribonuclease-A, lysozyme, trypsin and
pepsin, monitoring the protein concentration and the concentration of water found in the
organic phase. It was observed that the transfer of water is generally moderate (below
4%), whereas, under certain conditions, the protein is quantitatively transferred. This fact
demonstrated that the transfer of the protein into the micellar phase is not a passive
process, i.e. is not simply due to the fact that water is transferred and with it the protein.
The conclusion was that there is a thermodynamic driving force for the hydrophilic
protein to leave the aqueous environment and to transfer into the reverse micelles. In
other words, it seems that under certain conditions the protein-reverse micelle complex is
energetically favoured above the free protein and empty reverse micelles. Interactions can
be electrostatic, when surfactants with charged head groups are used, or hydrophobic
with the surfactant interface or the apolar solvent.
188 J. A. Asenjo and J. B. Chaudhuri
The fact that electrostatic interactions play an important role in the distribution of
proteins over reverse micellar and aqueous phase is shown by the dependence of the
aqueous phase pH and ionic strength.
The pH of the solution will affect the solubilisation characteristics of a protein prima-
rily in the way in which it modifies the charge distribution over the protein surface. With
increasing pH the protein becomes less positively charged until it reaches its isoelectric
point (PI). At pHs above the PI the protein will take on a net negative charge. If electro-
static interactions play a significant role in the solubilisation process, partition with
anionic surfactants should be possible only at pHs below the PI of the protein, where the
protein is positively charged and electrostatic attractions between the protein and the
surfactant head groups are favourable. At pHs above the PI, electrostatic repulsions
would inhibit protein solubilisation.
Goklen and Hatton (1987) have presented results on the effect of pH on solubilisation
of cytochrome-c, lysozyme, and ribonuclease-A, in AOT/isooctane reverse micelle solu-
tions. The results were presented as the percentage of the protein transferred from a
1 mg/ml aqueous protein solution to an equal volume of isooctane containing 50 mM of
the anionic surfactant AOT. A summary of their results is presented in Table 7.3.
Table 7.3. Effect of pH on solubilisation
Protein PI pH range of maximum
solubilisation
cy tochrome-c 10.6 5-10
ribonuclease-A 7.8 1-7
lysozyme 11.1 6-1 1
As anticipated, only at pHs lower than the pl was there any appreciable solubilisation
of a given protein, while above the PI the solubilisation appears to be totally suppressed.
However, at extremes of pH there is a drop in the degree of solubilisation of the proteins
due to protein denaturation, observed as precipitate formation at the interface (Chaudhuri
et al., 1993).
Luisi et al. (1979) used the quaternary ammonium salt methyl-trioctylammonium
chloride (TOMAC) for the transfer of a-chymotrypsin from water to cyclohexane. It was
found that the pH had to be reduced to values significantly below the PI (PI = 8.3) for
there to be any appreciable solubilisation. The solubilisation occurred only over a very
narrow pH range before decreasing rapidly again with further decreases in the pH of the
aqueous feed phase, accompanied by precipitation at the interface.
Similar results have been obtained by Dekker et al. (1986) for the enzyme a-amylase.
Significant solubilisation of the enzyme was observed over a narrow pH range in the
vicinity of 10-10.5 (PI = 5.1). In this pH range, all basic residues will be deprotonated
and the only charged residues being the carboxyl groups bearing a negative charge.
Innovative separation methods in bioprocessing 189
These results suggest that for higher molecular weight proteins a precise match be-
tween the surface charge densities of the protein and the reverse micelle is needed for the
transfer to occur.
The effect of the ionic strength of the aqueous phase is primarily to mediate the
electrostatic interactions between the protein surface and the surfactant headgroups. AS a
result of Debye screening, increases in the ionic strength of the protein feed solution can
be expected to reduce the interaction between the protein and surfactant headgroups,
hence decreasing the solubilisation of the protein. This has been studied by Goklen and
Hatton (1987). The transfer to the reverse micellar phase decreases for all proteins stud-
ied (cytochrome-c, lysozyme and ribonuclease-A) at increasing ionic strength, but the
point where this decrease starts depends on each particular protein. Such studies have
shown that protein transfer is influenced by electrostatic interactions between the protein
and surfactant headgroup, and that selectivity with respect to other proteins can be con-
trolled by manipulation of pH and ionic strength (Goklen and Hatton, 1987).
Other recent fundamental research in this field has focused on aspects of the kinetics
of protein partition and the structure of the microemulsions (Fletcher et al., 1987; Luisi et
al., 1988). Other work has indicated the considerable potential for the use of reverse
micelles in selective protein purification (Dekker et al., 1989). Woll et ul. (1989) have
shown that affinity ligands can be accommodated within the micelles to enhance
selectivity.
A study on the fractionation of intracellular proteins from Bakers’ yeast (Chaudhuri,
1991) proved to be complex, with problems of low protein yield, interfacial precipitation
and denaturation, which were not evident in the studies with model solutions. As a result
of disrupting the yeast cells, the intracellular contents (sugars, lipids, nucleic acid etc.)
will be present as ‘impurities’. It is not known if these non-protein components
contributed to the observed precipitation and denaturation, or if they can partition into the
microemulsion, thereby reducing protein capacity and selectivity.
There has been some activity in this field: for example, Schomaecker et al. (1988)
have observed the inactivation of a-chymotrypsin in an AOT/n-heptane system. How-
ever, they concluded that activity was lost as a result of enzyme autolysis as well as by
interaction with the microemulsion system.
Finally, the use of reverse micelles to isolate and refold pure denatured proteins has
been studied (Hagen et al., 1990a) as shown in Fig. 7.7. The results reported 100%
renaturation of denatured ribonuclease-A refolded in the reverse micelles. This method
proved unsuitable for the refolding of interferon-y because its hydrophobic nature causes
it to aggregate during the extraction process (Hagen et al., 1990b). There remain several
unanswered questions regarding the use of reverse micelles, for example, how applicable
is this method to other proteins, how will protein from inclusion bodies behave in this
system, and is any activity lost during the extraction process?
7.3.3 Perfluorocarbon affinity separations
Affinity chromatography exploits the natural, biospecific interactions that occur between
biological molecules. These interactions are very specific and because of this affinity
separation processes are very high resolution methods for the purification of proteins.
However, conventional affinity chromatography has its drawbacks. Namely the gel
190 J. A. Asenjo and J. B. Chaudhuri
Inclusion
body
F Refolded
Solubilised
protein
j. 4 u k protein
4hOc-O
4' k Reverse
7 n x micelle
Fig 7 7 Schematic diagram of protein refolding in reverse micelles
matrix used as the support for the affinity ligand is not very stable towards extremes of
pH which are found during sterilisation, and fouling of the columns may occur unless all
particulate material is removed prior to application on the column. This requires the use
of a solid-liquid separation operation increasing the process costs.
A development in affinity purification is the use of perfluorocarbons as supports for
affinity chromatography. Perfluorocarbons are synthetic molecules consisting of only
carbon and fluorine. These compounds are chemically and biologically inert and are
insoluble in both organic and aqueous solutions. However, the extreme hydrophobicity of
perfluorocarbons would lead to protein denaturation on contact. The perfluorocarbon can
be wetted in the presence of fluorosurfactants which adsorb to the perfluorocarbon By
attaching a triazine dye molecule to the fluorosurfactant an affinity ligand IS constructed.
The perfluorocarbon emulsion is formed by homogenising the perfluorocarbon in the
presence of the surfactant. The resulting emulsion droplets have diameters ranging from
10 to 37 pm (McCreath et al., 1992). The droplet size may be manipulated through the
emulsification conditions. Following emulsification the surfactant is cross-linked using
glutaraldehyde in the presence of HCl. The reduction of ligand leakage may be achieved
by derivatising poly(viny1 alcohol) PVA with fluoroalkyl groups and then using this to
coat the perfluorocarbon surface.
Perfluorocarbon emulsions have been utilised in an expanded bed configuration. The
perfluorocarbon emulsion must be used in an expanded bed to get round the problems of
droplet compression which would occur in a fixed bed column. The additional advantage
of this method is that the expanded bed arrangement allows any solid particles to flow
around the suspended emulsion droplets, thereby giving the potential for direct product
removal from fermentation or cell culture broths. The emulsion droplets are allowed to
settle in a column and are fluidised by the upward flow of buffer through the bed. At the
maximum flow rate this resulted in an expanded bed twice the height of the settled bed
(McCreath et al., 1992). By the use of the ligand CI Reactive Blue 4 such a system has
been used to adsorb human serum albumin (HSA) from plasma (McCreath et al., 1992).
The HSA was recovered at 87% yield and 91% purity. The purification factor was 1.44.
Innovative separation methods in bioprocessing 19 1
More recently this technique has been exploited in a continuous reactor for protein
purification. Current chromatographic practice results in batch protein purification with
the product being eluted and recovered at one stage in the process. Continuous separation
would enable the protein to be recovered continuously, and would be amenable to scale
up more easily than conventional chromatography. The emulsion reactor is based on
liquid-liquid contact between the perfluorocarbon emulsion and the protein solution. The
high density of the perfluorocarbon (1.8-2.1 g/ml, Stewart et al., 1992) results in fast
settling of the emulsion and aqueous phases and thus is suitable for a liquid-liquid
extraction process. The protein is adsorbed onto the affinity perfluorocarbon emulsion
which separates from the depleted aqueous phase under gravity. The loaded
perfluorocarbon emulsion is eluted with a buffer to recover the protein and then
re-equilibrated for further use. Continuous protein purification is carried out using a four-
chambered mixer-settler type configuration known as a perfluorocarbon emulsion reactor
for continuous affinity separations (PERCAS) (McCreath et al., 1993). Each chamber is
identical and consists of a mixing zone agitated by a turbine. The perfluorocarbon-
aqueous mixture passes over a weir and into a settling chamber, where the
perfluorocarbon emulsion settles to the bottom with the aqueous phase on top. Either
phase can then be pumped out either into the top of the next chamber or to waste. In the
first chamber the adsorption of the protein onto the emulsion takes place. After settling
the loaded emulsion is pumped into the second mixing chamber, where it is washed to
remove any trapped contaminants. The washed emulsion is pumped into the third
chamber, where the protein is eluted and subsequently recovered. The depleted emulsion
into the first chamber to adsorb more protein.
This process was tested using the adsorption of HSA from plasma using the dye ligand
CI Reactive Blue 2 attached to the perfluorocarbon emulsion described above (McCreath
et al., 1993). Protein binding was found to fit the Langmuir isotherm. The continuous
protein separation was controlled by varying the aqueous flow rates; the emulsion flow
rate was kept constant. The total protein recovery was 89% with HSA recovery at 81%.
The overall process yield was 71% with the HSA recovered at 91% purity (purification
factor of 1.52).
The perfluorocarbon emulsions behave as normal chromatographic materials when
operated in a fluidised bed. The protein adsorption capacities are comparable to conven-
tional matrices. Advantages of these materials are the fast adsorption and desorption
which arise as the droplets are non-porous. The advantages of the PERCAS system are
the continuous nature of the operation and the relative simplicity of the system - a mixer-
settler combined with gravity settling.
7.3.4 Liquid membrane separations
Liquid membrane extraction is a relatively new separation technology which has signifi-
cant potential for the selective separation and concentration of low molecular weight
chemicals produced by fermentation and used in the food-processing industries. Separa-
tion is achieved by the transport of the solute from a feed phase across a film of organic
solvent into a stripping phase. Examples of products successfully extracted using liquid
membranes include organic acids such as citric acid (Boey et al., 1987), lactic acid
is F""@ in20 3 fL\l?fih &aT[>her, Jyl!qwq i> is >T,3&,& 2nd ;rc.-q,i>iha..l&, ax,d s!qex, pup@
192
(Chaudhuri and Pyle, 1990), and amino acids such as L-phenylalanine (Itoh et al., 1990).
Currently, the recovery and purification processes for these species involve several steps
specific to individual manufacturers. There is, therefore, scope for the application of
liquid membranes which are generally one-step processes, and can simultaneously sepa-
rate and concentrate the solute. A schematic diagram of a liquid membrane process is
shown in Fig. 7.8.
The liquid membrane consists of the organic solvent which separates the two aqueous
phases (the feed and stripping phases), and contains a camer species to enhance both
selectivity and rates of extraction. Aliphatic diluents are generally preferred as the solvent
because of their lower solubility in water. In an ideal situation the solvent should have no
solubility in water to ensure that there is no aqueous phase contamination by trace
organics. There are two main configurations by which liquid membrane extraction can be
exploited, as discussed in the next two paragraphs.
A supported liquid membrane (SLM, Fig. 7.9) can be achieved by impregnating a
porous solid film with an organic solvent, which is held in place by capillary forces that
exist within the pores In order for the membrane pores to be effectively wetted, the
surface tension of the solvent must be less than the critical surface tension of the mem-
brane polymer. The membrane separates an aqueous phase, initially containing the re-
quired species, from an aqueous phase into which the solute is extracted, the stripping
phase. The solid supports used are generally microporous polymeric films, e.g.
Feed Liquid Stripping
phase membrane phase
phase
J. A. Asenjo and J. B. Chaudhuri
Solute - -
transfer
Fig 7 8 Schematic diagram of liquid membrane processes
Solid membrane
Stripping
phase
External
phase
+
Solute flux
Fig 7 9 Schematic diagram of supported liquid membrane (SLM)
Innovative separation methods in bioprocessing 193
polypropylene, polysulphone, or other hydrophobic materials. Typical dimensions are a
membrane thickness of 25-50 pm, with pore sizes of 0.02-1.0 pm.
An emulsion liquid membrane (ELM, Fig. 7.10) is formed by creating, under high
shear, a dispersion of the stripping phase within the organic solvent which forms a non-
porous film around the stripping phase droplets. The emulsion thus formed (stabilised by
a surfactant) is dispersed into the feed phase containing the solute, which is then trans-
ported into the stripping phase. Depending on the dispersion conditions the globule
diameter is 1-2 mm and the internal phase droplets are micron sized. The two aqueous
phases cannot physically contact each other and the solute is transported into the internal
phase droplets by diffusion through the stabilised solvent film. The use of a chemical
reagent in the stripping phase, which reacts with the extracted solute, prevents the solute
from diffusing back across the membrane phase. This strategy allows the removal of
virtually all of the solute from the feed solution, which makes emulsion liquid membrane
extraction very attractive for the recovery of solutes formed in low concentration. After
the extraction step the solute is recovered by allowing the emulsion and feed phases to
separate, by settling under gravity, then removing the emulsion and breaking it to release
the separated solute from the membrane phase components. Electrostatic splitting is
generally used for de-emulsification as the membrane components can be recycled for
further use.
ELM systems give rise to very fast extraction kinetics and allow the use of conven-
tional liquid-liquid extraction equipment; they are also prone to emulsion
swelling, which gives rise to dilution and instability problems, The necessity to make and
break an emulsion does not arise with the SLM system; however, this configuration has
slower kinetics, and loss of the membrane phase may occur. In summary, liquid
membrane processes offer high separation factors, low capital and operating costs, a
lower solvent inventory than solvent extraction, ease of scale-up and the possibility of
continuous operation.
The key to selectivity in liquid membrane extraction is the use of a carrier species
incorporated in the organic solvent to increase the solute solubility: by introducing a
0 External
w< 0
c>
0
0 Membrane
phase 0 **@
0"yp
Internal phase
0 droplets
0
Fig. 7.10. Schematic diagram of emulsion liquid membrane (ELM).
194
‘carrier’ molecule into the membrane phase, the solute solubility is increased by the
reversible formation of a membrane-soluble carrier-solute complex. This results in faster
mass transfer rates, and selectivity is introduced into the extraction as the carrier-solute
reaction can be selective. This is known as facilitated transport. The use of a carrier
enhances selectivity by the formation of a reversible complex between the carrier and the
solute, which is only soluble in the organic solvent. This is particularly effective for the
recovery of charged solutes which may be poorly soluble in the organic solvent. Many of
the carriers so far employed in liquid membrane processes are extractants used in
conventional liquid-liquid extraction, e.g. secondary and tertiary amines, and
phosphorus-containing extractants.
There have been several applications of liquid membrane extraction to biotechnologi-
cal separations. The recovery of citric acid has been studied by Boey et al. (1987). The
liquid membrane consisted of Alamine 336 and Span 80 dissolved in Shellsol A. Sodium
carbonate was used as the internal phase reagent. This work looked at the batch extrac-
tion of both model and real fermentation broths. The results show that very fast extraction
of citric acid can be achieved: over 80% of a 5% (w/v) citric acid solution was removed
in under 5 min. Significant emulsion swelling was also observed in this study; the volume
of the internal phase was more than doubled. The recovery of citric acid in supported
liquid membranes has also been reported (Sirman et al., 1990).
Lactic acid extraction has also been achieved using emulsion liquid membranes
(Chaudhuri and Pyle, 1992a, b). The tertiary amine Alamine 336 was used as the carrier
species. Extraction yields were found to be low when recovery was attempted from real
fermentation broth. This was a result of competition for the carrier by ‘impurities’ in the
broth. In this study significant swelling was observed; however, it was also possible to
eliminate the effects of swelling by adjustment of the initial osmotic pressure difference.
Scholler et al. (1993) have found that recovery of lactic acid from homogeneous
‘model’ solutions resulted in fast extraction with significant product concentration, whilst
extractions attempted from Lactobacillus delbreuckii fermentation broth exhibited poor
extraction kinetics and yields. This was believed to be a result of low carrier selectivity
for the solute of interest and competition by other compounds present.
Work on the extraction of phenylalanine using emulsion systems has been carried out
(Itoh et al., 1990). In common with other studies, work has focused on the recovery of the
solute from a ‘model’ fermentation broth. In this study the carrier used was Aliquat 336
(tricaprylmethylammonium chloride), in Solvent 1 OON, stabilised by the surfactant
Paranox 100. In common with other work, significant emulsion swelling was observed. It
was found that swelling increased with the concentration of the internal phase reagent.
The nucleotides adenosine and deoxyadenosine have been transported through a
chloroform liquid membrane using lipophilic camers. Adenosine mono- and diphosphate
(AMP and ADP) have also been transported through chloroform using a lipophilic
diammonium salt of diazobicyclooctane (Pellegrino and Noble, 1990).
Chiral separation, or the resolution of optical isomers has been attempted for the
transport of amino acids with 4-28% enantioselectivity. Chiral transport of sodium
mandelate has been used on the optically active carrier N-( 1 -napthyl)methyl- 7a 1-
methylbenzylamine dissolved in chloroform. The choice of anion was found to influence
chiral selectivity. Separation of racemic N-(3,5-dinitrobenzoyl) 7a 1-amino acid
J. A. Asenjo and J. B. Chaudhuri
Innovative separation methods in bioprocessing 195
derivatives was achieved using (S)-N-( 1 -napthyl)-leucine octadecyl ester as the carrier in
dodecane (Pellegrino and Noble, 1990). Racemic amine salts have been separated into
their optical isomers by complexation with an optically active macrocyclic ether and
transported from one aqueous phase, through chloroform, into a second aqueous phase.
The application of liquid membranes as a separation technique has mainly been in the
field of hydrometallurgy, but more recently in biotechnology. Some of these applications
have been described above. It is likely that more applications will become apparent
through advances in carrier chemistry.
7.4 SOLID-BASED SEPARATIONS
Various new and unconventional technologies for biological molecules are presently
being developed. These include systems that facilitate the handling of r.aterials by the
use of membrane, fluidised bed and novel chromatographic matrix technologies,
improving separation specificity and efficiency by using metal, dye and other ligands and
also by developing techniques used very efficiently at the analytical scale such as
isoelectric focusing now at preparative scale.
7.4.1 Adsorption systems: expanded bed adsorption
Expanded bed adsorption or fluidised bed adsorption is a new technique that in one step
accomplishes removal of whole cells and cell debris, concentration and initial purification
of the target protein. It is based on the design of an adsorbent whose density enables the
formation of a stable expanded or fluidised bed (two- or threefold expansion) that allows
the establishment of a plug-flow concentration profile across the column. Binding
capacities for proteins are similar to those obtained in fixed bed adsorption and linear
velocities can be higher than 300 cm/h (Janson and Arve, 1993). Pharmacia has recently
started to commercialise expanded bed adsorption under the trade name Streamline. This
operation allows the fusing of two or three operations, namely clarification, concentration
and capture, into one operation. At the outset, the adsorbent is sedimented in the column.
The bed is then expanded by pumping buffer upward through the column. Once the bed is
expanded, the buffer is replaced with crude feed which contains the products, cells, cell
debris, contaminants and other particulates. After the product is adsorbed, loosely bound
material is flushed with a buffer wash and the direction of the liquid flow is reversed with
the change to elution buffer and the product is eluted as in packed bed chromatography. If
the product is extracellular the process stream is the crude fermentation broth and if it is
intracellular the stream is the crude homogenate.
The adsorption behaviour of a fluidised bed has been investigated in detail (Draeger
and Chase, 1990). A frontal analysis during adsorption to investigate the shape of the
breakthrough curve is shown in Fig. 7.1 1. The shapes of the two curves for the adsorption
of BSA onto Q-Sepharose FF are very similar, suggesting that the fluidised bed system is
behaving in much the same way as the fixed bed system. The apparent maximum adsorp-
tion capacity calculated from both runs was 80 g/l, which is consistent with the values
from the batch isotherm experiments. A theoretical model for predicting adsorption per-
formance in a fixed bed was used to fit theoretical curves to the data from both a fluidised
or expanded bed and fixed bed systems (Draeger and Chase, 1990). The model
196 J. A. Asenjo and J. B. Chaudhuri
predicted the performance of the system very accurately, indicating again the similarity
between the fluidised and fixed bed systems.
Fig. 7.11. Fixed and fluidised bed adsorption of BSA to Q-Sepharose fast flow in 0.01 M Tris
buffer pH 7.0 (0 fixed bed, o fluidised bed).
7.4.2 Continuous adsorption recycle extraction
Continuous adsorption recycle extraction was recently adapted for biotechnological ap-
plications as a downstream process for efficient separation of proteins from crude
feedstocks to be carried out continuously (Pungor et at., 1987; Gordon et at., 1990).
Modem biotechnology involves relatively small-scale processes that favour batch proce-
dures, usually following the concept that a batch is well defined (from beginning to end)
and that this will pose fewer problems with regulatory approval for therapeutic use.
However, it is well known that in the chemical and biochemical process industries batch-
to-batch variations can be substantial. It is also well known that a continuous process can
operate under steady-state conditions minimising batch-to-batch variations. In addition, a
continuous process is much more appropriate for implementing a control strategy in order
to maintain constant conditions more readily (Rodrigues et at., 1992).
Continuous adsorption recycle extraction, also called CARE, is based on the continu-
ous protein adsorption to solid phase supports and features two well-mixed reactors with
solids recycle, as shown in Fig. 7.12. The adsorbing stage takes place in the first reactor
(adsorber) where the liquid feed is contacted with the adsorbent; desorption of the protein
to be purified is obtained in the second reactor (desorber) by maintaining appropriate
conditions and by the action of a suitable eluting solvent. The adsorbent beads are
recycled to the adsorber reactor while the product is continuously removed. The system' s
performance is determined by the nature of the adsorbent and the feed material, flow
rates and reactor volume and the conditions of the adsorption and de sorption stages. An
accurate mathematical model has recently been developed which can be used for investi-
gating optimisation criteria, but also is important for system design and process control
(Rodrigues et at., 1992).
Innovative separation methods in bioprocessing 197
Eluent
, F;ue, Desorber
Adsorber
recycle
Waste Product
Fig. 7.12. Continuous adsorption recycle extraction (CAFE) process (see text for key to symbols).
During operation, the feed composition and flow rate may experience uncontrolled
variations which may upset the steady state. This is also true for adsorbent binding
capacity and even recycle flow rates and composition. Hence, appropriate control
schemes are needed to meet product specifications and achieve operational stability in the
face of potential external and internal disturbances.
As a first approximation, the two reactors (adsorber and desorber) are assumed to be
perfectly mixed (Pungor et al., 1987). The adsorption process is regarded as a reversible
second-order reaction, whereas the desorption stage is modelled as a first-order irrevers-
ible reaction scheme (Rodrigues et al., 1992):
kl
k2
adsorption: A+B + AB
k3
desorption: AB --+ A + B
where A, B and AB are the target protein, the adsorbent and the adsorbed protein,
respectively. The rate constants, k,-,, represent not only the intrinsic adsorption and
desorption kinetics, but also include contributions from both external and internal mass
transfer resistances (Sherwood et al., 1975). Thus, these effects are lumped into a single
coefficient which can be determined experimentally (Chase, 1984). This approach makes
the mathematical formulation more tractable at the expense of a less rigorous physical
description. On the other hand, experimental evidence suggests that in such a system
external mass transfer limitations can be neglected without loss of accuracy (Pungor et
al., 1987) and, in the case of high molecular weight proteins, these are mostly adsorbed in
active sites located at or near the gel’s surface since molecules attached to those sites
block the way to diffusion into the gel. Any thermal effect associated with adsorption and
desorption is neglected and the operation is carried out under isothermal conditions. The
ordinary differential equations system can be easily solved using the Runge-Kutta fourth-
order algorithm. The steady-state solution is obtained by setting the accumulation terms
(Le. time derivatives) equal to zero, giving rise to a set of non-linear algebraic equations
which can be solved analytically. The steady-state model provides a valuable tool in the
formulation of control schemes.
Table 7.4 shows the parametric sensitivity of steady state outputs, when each input
varies 20% around the baseline value. The sensitivity index S was calculated as the ratio
between the percentage variation of the output and the percentage variation of the input,
198 J. A. Asenjo and J. B. Chaudhuri
Table 7.4. Steady state sensitivity analysis (sensitivity index, S)
output
Input Cl c2 41 42 C2/CO
F1 0.7 0.3 0.3
0.3 0.2
4m -0.7 0.6
0.6 0.6 0.1
kl -0.3 0.2 0.2 0.2 0.2
k2 0.3 -0.2 -0.2 -0.2 -0.2
k3 0.0 0.0 0.0 0.8 0.0
CO 1.2 0.4 0.4 0.4 0.4
F2 -0.1 -1 .o -0.1 -0.1 -0.6
FR -0.6 0.5 -0.3 0.7 0.4
q, = bound protein concentration in adsorber
q2 = bound protein concentration in desorber
qm = maximum adsorption capacity of gel
All other variables and parameters are defined in Fig. 7.13 and in text.
and represents the corresponding steady state gain. It can be seen that all outputs are
strongly affected by the feed protein concentration (CO), the gel recycle rate (FR), the gel
maximum adsorption capacity (qm) and, to a lesser extent, by the feed flowrate (F1). The
eluent flowrate (F2) has a strong influence on the product stream concentration (C2) and
the protein recovery ratio (C2/C0). The kinetic parameters (kl-3) do not seem to have a
strong influence on performance under the baseline conditions; in the case of k3, 42 is the
only output showing some sensitivity. Figure 7.13 shows the time response of protein
concentration in the product stream after step changes in feed, eluent and gel recycle flow
rates.
-
O 1.2
G
8 1.0
cu
2 0.8
G 20% increase in:
I 0.6
2
2 0.4
- a, FR
.e 0.2
00 10 20 30 40
.E 0
0
cu
c
m
c a,
Time (hours)
Fig. 7.13. Open-loop response of dimensionless production concentration to a 20% step change
in input flow rates (Rodrigues er al., 1992). F1 = feed flowrate; F2 = eluent flowrate;
FR = recycle flowrate; CO, C1, C2 = protein concentration in streams shown.
Innovative separation methods in bioprocessing 199
In conclusion, a simple and effective process control scheme to regulate the
continuous adsorption recycle extraction of proteins has been investigated. A dynamic
mathematical model, based on unsteady mass balances, was derived and used in the study
of the steady state and dynamic behaviour of the system, and to assist in the synthesis of
appropriate control schemes. Sensitivity analysis showed that performance was strongly
affected by the feed flow rate and feed protein concentration. The eluent flow rate was
shown to be the most suitable manipulated variable. The protein concentration in the
product stream was successfully controlled using a conventional PI (proportional
integral) feedback controller, when measurement time delays are not significant (less than
50% of sampling interval). Model-based feedforward compensation further improved
performance in those cases where the main disturbances can be measured on-line in the
feed stream.
The control philosophy developed can be extended to other separation processes such
as continuous two-phase aqueous separations featuring continuous liquid-liquid extrac-
tion (Asenjo et al., 1991; Hustedt et al., 1988).
7.4.3 Membrane chromatography
Chromatography is a well-used and well-understood method of purifying proteins. How-
ever, limitations in this process arise from mass transfer effects that occur in the beads.
As a result of the relatively large diameter and porosity of the beads, diffusion times into
the beads may be long, which results in broad peaks. In order to shorten the process time
new efficient adsorbents are sought. One option was to reduce the bead size to minimise
the effects of diffusive transport into the pores. This option gives rise to high pressure
drops across the bed. A recent development is that of perfusion chromatography in which
the adsorbents have many larger pores allowing greater flow velocities without
compromising the resolution (Afeyan and Regnier, 1990). An alternative approach is to
use membrane matrices as adsorbents. These have favourable characteristics such as high
surface area, high solute throughput, mechanical strength, a highly porous structure and
low pressure drops (Fig. 7.14). A potential advantage of membrane chromatography is its
use in unclarified broth streams. There are also various process configurations which may
be exploited. Membranes that are going to be used as adsorbents must have a high
hydrophilicity, low non-specific protein adsorption, a uniform structure, a narrow pore
size distribution, chemical and mechanical resistance and the ability to allow ligand
coupling (Briefs and Kula, 1992). Ligand coupling may limit the modes of adsorption
and types of interaction that may be suitable for membrane chromatography. Dye affinity
ligands have been successfully coupled directly to the surface of nylon membranes;
however, it was not possible to create an anion exchange membrane by this method and
the use of a dextran spacer was required (Briefs and Kula, 1992).
Affinity membranes have been created by the coupling of triazine dyes to the mem-
brane surface. These dyes have been successfully exploited as agents that mimic enzyme
cofactors and other prosthetic groups, and thus may be used as affinity ligands. The
technique of dye-ligand chromatography has been developed as a high-resolution purifi-
cation technique.
The membranes that have been used are nylon-based (Ulitpore, Loprodyne,
Immunodyne), and the dyes have been coupled to the membranes either directly or using
200 J. A. Asenjo and J. B. Chaudhuri
Flow distributor
Membrane stack
Gaskets
Flow distributor
Fig. 7.14. Membrane chromatography system (Briefs and Kula, 1992).
1,6-diaminohexane or polyethylenimine spacers (Champluvier and Kula, 1990). These
membranes are flat circular discs used in the laboratory for sample preparation.
Affinity membranes have been tested by batch adsorption and by membrane
chromatography. In batch adsorption experiments pieces of dye-membrane are shaken
with protein solutions in a test tube. In the membrane chromatography mode the dye-
membrane discs were encased in a filter holder and protein solutions and buffers intro-
duced with a syringe. Alternatively, up to 10 filter housings were stacked and connected
to an FPLC apparatus (Pharmacia, Sweden) (Champluvier and Kula, 1991).
Experiments were carried out where the membrane-ligand system was compared with
a bead-ligand system (Sepharose CL-4B) (Champluvier and Kula, 1990). Adsorption of
malate dehydrogenase into Cibacron blue dye was performed. The loading of the protein
was similar whatever the carrier. However, higher desorption of the protein was obtained
with the beads. Further studies with the proteins adenylate kinase and glucose-6-
phosphate dehydrogenase and beads of Sepharose CL4B with yellow and blue dyes were
carried out. The volumetric capacity of the membranes was similar to that of the gel
beads. However, desorption was poor with respect to the Sepharose. The effect of the
spacer was found to be very important. The dye concentration attached to the membrane
was found to be five to ten times greater when polyethylenimine spacers were used as
compared to directly binding the ligand to the membrane (Champluvier and Kula, 1991).
Similarly, adsorption of lysozyme was found to be significantly higher onto dye-
membranes rather than the non-specific binding to the undyed membrane.
Affinity membrane chromatography has been tested with the purification of glucose-6-
phosphate dehydrogenase (G6PDH) from Saccharomyces cerevisiae as a case study
(Champluvier and Kula, 1992). The ligand used was Cibacron blue F3GA bound to a
Sartorius membrane with a thickness of 210,um and pore size of 0.45 gm. The
conclusions of this study were that the particles should be removed prior to adsorption to
Innovative separation methods in bioprocessing 201
assure fast operation and long-term performance. Adsorption of dilute enzyme solution
was found to be very fast and results from the convective flow through the adsorbent and
the fewer short diffusion pathways. Less than 25% of the bulk protein was adsorbed
during the loading step, showing good selectivity. Desorption with KC1 and PEG was
slow and gave a trailing peak. This was improved dramatically by eluting with the
cofactor NADP. Scale-up of this process to a membrane with a cross-sectional area 40
times greater was achieved. The eluted protein concentrations were similar as were the
purification factors and yields. Scale-up may also be performed by stacking membranes,
provided unacceptable pressure drops do not occur at the desired flow rates.
Other enzymes that have been purified by affinity membrane chromatography are
pyruvate decarboxylase (PDC) from Zyrnornonas rnobilis and formate dehydrogenase
(FDH) from Candida bodinii (Briefs and Kula, 1992). Both enzymes are amenable to
purification by dye-ligand chromatography; PDC was recovered using Procion yellow
HE-4R and FDH by Procion red HE-3R. PDC was purified with an anionic membrane
system and two gel matrices; DEAE-Sephacel and Mono-Q. It was found that the quality
of resolution obtained by the membrane was better than the DEAE-Sephacel and nearly
as good as the Mono-Q column. The cycle times for purification were shortest with the
membrane, 1 minute as compared with 21 minutes for the Mono-Q column and 420
minutes for the DEAE-Sephacel system. As part of this study membrane chromatography
was mathematically modelled. It was found that the kinetic effects leading to peak
broadening can be neglected at flow rates less than 5 cm/min for affinity adsorption, and
less than 20 cm/min for ion exchange. Secondly, the size of the pores has only a minor
effect on resolution.
In conclusion it is evident that the high flow rate and high efficiency of capture
characteristics of membrane chromatography are well suited for processing large volumes
of dilute protein solutions. Choice of membrane will be very important. It was found that
the available area for binding was not dramatically reduced when membranes with larger
pore sizes were used. Pores with diameters of 1.2-3.0 pm are suitable for low pressure,
high capacity and good resolution.
7.4.4 Chromatographic and adsorption materials
Perj5usion chromatography represents an interesting approach to overcoming speed and
throughput limitations in chromatography caused by diffusive mass transport. The pack-
ing used for perfusion chromatography contains two classes of pores: throughpores (see
Fig. 7.15) to allow convective flow through the particles and smaller, difSuusive pores
lining the throughpores to provide high adsorption surface area. Conventional materials
contain only diffusive pores and thus solute transport to the binding sites on the medium
is limited by the slow diffusion process.
In perfusive media, when the mobile phase reaches a sufficiently high velocity, con-
vection in the throughpores takes place. In this regime, the combined rapid convective
transport and ultra short path length in the diffusive pores make resolution and dynamic
loading capacity virtually independent of flow rate with linear velocities, ca. ten times or
even higher than with conventional materials.
Inverted matrix chromatography uses a resin matrix where the resin is the continuous
phase in the matrix and the void spaces are distributed. This is in contrast to the normal
202 J. A. Asenjo and J. B. Chaudhuri
Fig. 7.15. Perfusion chromatography.
resins which consist of discrete particles. Very low pressure drops are possible and their
use in direct broth extraction is presently being investigated (Howell et al., 1993).
Affinity chromatography has been used for a number of years both at the laboratory as
well as at the large scale. Most early industrial processes for therapeutic proteins included
one or more affinity steps. This corresponds to a very high-resolution purification even in
the absence of physicochemical knowledge about the main contaminants (see Table 7.1).
These, however, relied in many cases on very large ligand molecules (e& antibodies) and
were usually exorbitantly expensive at either the large scale or even the laboratory scale.
At present there is a tendency to use less expensive, more ‘generic’ ligands. Important
advances in this area have been obtained in the last few years which include the synthesis
and use of specific dyes and also of metal ions. This technique is also known as immobi-
lised metal affinity chromatography (IMAC).
7.5 OTHER DEVELOPMENTS
7.5.1 Electrically enhanced separations
Electrophoresis is capable of resolving biological molecules on the basis of differences in
their molecular weights, isoelectric points and mobilities in an electric field with a very
high resolution. At the laboratory scale this includes some of the most powerful tech-
niques available for the purification of biologically active molecules. Isoelectric focusing
(IEF) and recycle isoelectric focusing (RIEF) are used in the laboratory as efficient
preparative techniques to fractionate complex protein mixtures. The main advantage is
the very high resolution; however, removal of ampholites and the low protein solubility at
its isoelectric point are disadvantages for large-scale use. Other techniques that have been
at least partially scaled up are recycle continuous flow electrophoresis (RCFE) and coun-
teracting chromatographic electrophoresis (CACE) (Asenjo, 1990).
As electrophoresis is seldom employed at the commercial scale primarily due to the
convective mixing problems caused by ohmic heating it has recently been shown that an
aqueous two-phase system can be used as a medium for electrophoretic transport and that
the liquid-liquid interphase provides stability against convection and facilitates product
recovery. The method combines the scalability of aqueous two-phase systems with the
selectivity of electrophoresis.
Electrically enhanced membrane processes are presently being investigated as poten-
tially useful techniques both for resolution of protein mixtures and for the separation of
disrupted microbial cells (Brors et al., 1993; White et al., 1993).
Innovative separation methods in bioprocessing 203
7.5.2 Genetic approaches to protein purification
Recombinant DNA techniques have also allowed the development of highly specific
separations. This has been carried out by adding a specific tail to the target protein by
genetic engineering techniques (e.g. the zz domain of protein A) which binds a specific
ligand (e.g. IgG or a more inexpensive ligand such as metal ion). After purification of the
protein the tail is removed usually by enzymatic cleavage.
7.5.3 Purification of intracellular proteins
A large number of proteins synthesised in Escherichia coli and yeast are intracellular.
These include intracellular protein particles such as recombinant vaccines. Hence the first
step in their separation consists of their extraction from the contaminant cell material.
Techniques for selectively carrying out the solubilisation or extraction of the protein are
presently being developed (Fig. 7.16). These include the use of solvents (e.g. toluene),
chelating agents (e.g. EDTA), detergents (e.g. Triton X-100) and chaotropic agents
(guanidine and urea). Also, the use of a pure lytic glucanase to selectively release
recombinant 60 nm protein particles (virus-like particles or VLPs) from yeast has
recently been reported (Asenjo et al., 1993). The VLPs have been used in trials to
manufacture an AIDS vaccine already for 3 years. When using the crude lytic complex
which in addition to glucanase contained some lytic protease, the protease components of
the complex were found to degrade the VLPs. The purified glucanase enzymes from
these complexes produced cell lysis without degradation of the VLPs. The Oerskovia
lytic glucanase enzyme released the recombinant protein particles selectively as it only
produced ca. 17% cell lysis compared to the use of the crude lytic enzyme preparation
(with lytic protease). This selectivity, which results in the release of the recombinant
particles with only a fraction of contaminating proteins, represents an improvement over
presently used mechanical or enzymatic cell disruption processes. This pure lytic
glucanase is presently being cloned in E. coli and Bacillus (Ferrer et al., 1993) strains
used for commercial production of extracellular enzymes. The large-scale availability of
)
/
1A u R
Mechanically disrupted cell Permeabilized cell
Fig 7 16 Conceptual comparison of mechanical disruption and permeabilisation for obtaining
intracellular product release
204
an inexpensive lytic p (1-3) glucanase will find use in improved processes for selective
recovery of intracellular proteins from yeast.
J. A. Asenjo and J. B. Chaudhuri
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