Modification of basic bioreators.pdf sff

PART 3
Engineering Principles for Bioprocesses
9
Operating Considerations
for Bioreactors
for Suspension
and Immobilized Cultures
So far we have discussed what cells are, how they work, and how to describe their growth
in simple reactors. We now begin our discussion of how to use these cells in processes.
We will explore some more complicated reactor strategies and why they might be consid-
ered for use in real processes. Chapter 10 will give more details on reactor design, and
Chapter II will detail how to recover products from these reactors. These chapters should
give the reader an understanding of how real bioprocesses can be assembled.
An important decision for constructing any process concems the configuration the
reactor system should take. The choice of reactor and operating strategy determines prod-
uct concentration, number and types of impurities, degree of substrate conversion, yields,
and whether sustainable, reliable performance can be achieved. Unlike many traditional
chemical processes, the reactor section represents a very major component (usually
:> 50%) of the total capital expenditures. Choices at the reactor level and of the biocatalyst
determine the difficulty of the separation. Thus, our choice of reactor must be made in the
context of the total process: biocatalyst, reactor, and separation and purification train.
245

247
(9.7)
X
rc,opl = ln -.!!L +IlmtI
rb Xo
Choosing the Cultivation Method
The ratio for rates of biomass formation is
Sec. 9.2
Most commercial fermentations operate with X,jXo '" 10 to 20. Thus, we would ex-
pect continuous systems to always have a significant productivity advantage for primary
products. For example, an E. coli fermentation with X,jXo = 20, ti = 5 h, and !lm = 1.0 h-l
would yield rc.op/rb = 8.
Based on this productivity advantage we might be surprised to leam that most com-
mercial bioprocesses are batch systems. Why? There are several answers.
The first is that eq. 9.7 applies only to growth-associated products. Many secondary
products are not made by growing cells; growth represses product formation. Under such
circumstances, product is made only at very low dilution rates, far below those values op-
timal for biomass formation. For secondary products, the productivity in a batch reactor
may significantly exceed that in a simple chemostat.
Another primary reason for the choice of batch systems over chemostats is genetic
instability. The biocatalyst in most bioprocesses has undergone extensive selection. These
highly "bred" organisms often grow less wel1 than the parental strain. A chemostat im-
poses strong selection pressure for the most rapidly growing cello Back-mutation from the
productive specialized strain to one similar to the less productive parental strain (Le., a re-
vertant) is always present. In the chemostat the less productive variant will become domi-
nant, decreasing productivity. In the batch culture the number of generations available
« 25 from slant cultures to a commercial-scale fermenter) for the revertant cell to out-
grow the more productive strain is limited. Cel1s at the end of the batch are not reused.
These considerations of genetic stability are very important for cells with recombinant
DNA and are discussed in detail in Chapter 14.
Another consideration is operability and reliability. Batch cultures can suffer great
variability from one run to another. Variations in product quality and concentration create
problems in downstream processing and are undesirable. However, long-term continuous
culture can be problematic; pumps may break, control1ers may fail, and so on. Mainte-
nance of sterility (absence of detectable foreign organisms) can be very difficult to
achieve for periods of months, and the consequences of a loss of sterility are more severe
than with batch culture.
One other factor determining reactor choice is market economics. A continuous sys-
tem forms the basis of a dedicated processing system--dedicated to a single product.
Many fermentation products are required in small arnounts, and demand is difficult to
project. Batch systems provide much greater flexibility. The same reactor can be used for
two months to make product A and then for the next three for product B and the rest of the
year for product C.
Most bioprocesses are based on batch reactors. Continuous systems are used to
make single-cell protein (SCP), and modified forms of continuous culture are used in
Waste treatrnent, in ethanol production, and for some other large-volume, growth-
associated products such as latic acid.
Let us consider some modifications to these reactor modes.
rc.opt = DoptXopt = IlmYX/SSO
Bioreactors for Suspension and Immobilized Cultures
246
Xopt = Yx/s{So + Ks -.[Ks(So +Ks)}
Thus, the best productivity that could be expected from a chemostat where Monod ki
ics apply is Dop' . Xopt, or
DoptXop, = YxlsllJI- J Ks 1[So +Ks -~Ks(So+Ks))
L Ks +So
Under normal circumstances 50 » Ks, so the rate of chemostat biomass production,
approximately
9.2. CHOOSINGTHE CULTIVATlON METHOD
One of the first decisions is whether to use a batch or continuous cultivation scheme.
Although a simple batch and continuous-flow stirred-tank reactor (CFSTR) represent ex-
tremes (we wiU soon leam about other reactors with intermediate characteristics), consid-
eration of these two extreme altematives will clarify some important issues in reactor
selection.
First, we can consider productivity. The simplest case is for the production of cel1
mass or a primary product. For a batch reactor, four distinct phases are present: lag phase,
exponential growth phase, harvestí~g, and preparation for a new batch (e.g., cleaning,
sterilizing, and filling). Let us define tI as the sum of the times required for the lag phase,
harvesting, and preparation. The value for ti will vary with size of the equipment and the
nature of the fermentation but is normally in the range of several hours (3 to 10 h). Thus,
the total time to complete a batch cycle (tc) is
I I Xm
t =- n-+tI
c Ilm Xo
where Xm is the maximal attainable cel1 concentration and Xo is the cell concentration
inoculation.
The total amount of cell mass produced comes from knowing the total amount at
growth-extent-limiting nutrient present and its yield coefficient:
Xm - Xo = Yx/sSo
The rate of cell mass production in one batch cycle (rb) is
Yx/sSo
rb =
(l/llm)ln(Xm/XO)+tI
As discussed in Chapter 6, the maximum productivity of a chemostat is found by differe
tiating DX with respect to D and setting dDX/dD to zero. The value for D optimal wb
simple Monod kinetics apply is given by eq. 6.83, and the corresponding X can be det,
mined to be

9.3.1. Chemostat with Recycle
Il X dS
FSo+aFS - V~ - (1+a)FS = V-
Yx/s dt
249
(9.13)
Figure9.2. Comparison
ofbiomass
con-
centrations
andoutputratesinsteadystates
ofchemostat
cultures
withandwithout
recy-
cle.Symbols:
XI = biomassconcentration
in
chemostal
withoul
recycle;
X2 = biomass
concentration
inchemostal
culturewithre-
cycle;R I= biomass
outpulrateperunilvol-
umewithoul
recycle;
Rz = biomass
oulput
raleofchemostal
withrecycle;
/lm = 1.00
hol; Sr= 2.0 gI1; KS= 0.010 gI1; YXIS =
0.5 g/g;concentration
faclor.C= 2.0; and
recyclerate,a = 0.5.
s= K,D(l+a-aC)
Ilm -D(I+a-aC)
0.6 0.8 1.0 1.2 1.4 1.6 1.8 2.0
Dilulion rale (h-')
X2
Modifying Batch and Continuous Reactors
At steady state, dS/dt = Oand
XI = Y~s [s _ K,D(I+a-aC) ] (9.14)
(l+a-aC) o Ilm -D(I+a-aC)
Effluent cen concentrations and productivities in a chemostat with and without cen
recycle are compared in Fig. 9.2. cen concentrations and productivities are higher with
cen recycle, resulting in higher rates of substrate consumption. Systems with cen recycle
are used extensively in waste treatment and are finding increasing use in ethanol produc-
tion. The application of cell recycle reactors in waste treatment is detailed in Chapter 16.
The equations differ from the case above due to the inclusion of a term for endogenous
metabolism (Le., kd). The basic concept of operation at flows above the "washout" rate ap-
plies when kd *- O.
Example9.1
ln a chemostat with cell recycle, as shownin Fig. 9.1, the feed flowrate and culture volumes
are F = 100mlJhand V = 1000ml, respectively.The systemis operatedunder glucoselimita-
tion, and the yield coefficient, Yts. is 0.5 gdw cells/g substrate.Glucose concentrationin the
D M
X1=llg Yx/s(So-S (9.11)
Substitution of eq. 9.9 when kd = Ointo eq. 9.11 yields
yM (S -S)
X = x/s o (9.12)
I (l+a-aC)
Therefore, the steady-state cen concentration in a chemostat is increased by a factor
of 1/(1 + a - aC) by cen recycle. The substrate concentration in the effluent is deter-
mined from eq. 9.9 and the Monod eq. 6.30, where endogenous metabolism is neglected,
andis
Then eq. 9.12 becomes
Figure9.1. Chemostat
with
Thecellseparator
couldbea
tank,a centrifuge.
oramicro:
device.
eX1
Ilne. = (l + a - aC)D = [1+ a(l- C)]D
Bioreaetors for Suspension and Immobilized CultureS
dXI
FXo+aFCXI-(l +a)FXI + Vil
oe' XI = V-dt
aF
248
where a is the recycle ratio based on volumetric flow rates, C is the concentration fac'
or ratio of cen concentration in the cen recycle stream to the cen concentration in the
aetor effluent, F is nutrient flow rate, V is culture volume, Xo and XI are cen concen
tions in feed and recycle streams, and X2 is cen concentration in effluent from the
separator.
At steady state, and if dX/dt = O and Xo= O (that is, sterile feed); then eq. 9.8
comes
Since C> 1 and a(1 _ C) < O,then Iloe' < D. That is, a chemostat can be operated at
tion rates higher than the specijic growth rate when cell recycle is used.
A material balance for growth-limiting substrate around the fermenter yields
Microbial conversions are autocatalytic, and the rate of conversion increases with cell
concentration. To keep the cell concentration higher than the normal steady-state level in a
chemostat, cells in the effluent can be recycled back to the reactor. cen recycle increases
the rate of conversion (or productivity) and also increases the stability of some systems
(e.g., waste-water treatment) by minimizing the effects of process perturbation. cens in
the effluent stream are either centrifuged, filtered, or settled in a conical tank for recy-
cling. .'
Consider the chemostat system with cen recycle as depicted in Fig. 9.1. A materill1
balance on cen (biomass) concentration around the fermenter yields the following
equation:
9.3. MODIFYING BATCH AND CONTlNUOUS REACTORS

251
(9.15)
(9.16)
(9.17)
Figun 9.3. 1vo-stage chemostat system.
F2
X2
S2
SI = Ks~
Ilm -11
Xl =~s(So-~)
S~
'F'
,
I
I
•
V2
dX2
FXI - FX2 +112 VzX2 = Vz-
dt
S2
~
F
S,
XI
Modifying Batch and Continuous Reaetors
v,
F
So
The biomass balance for the second stage yields
stage, the inducer is added and large quantities of product are made. Cells defective in
product synthesis should not overtake the culture (at least not completely), because fresh
genetieally unaltered cells are being continuously fed to the reactor. Thus, the two-stage
system can allow the stable continuous production of the target protein when it would be
impossible in a simple chemostat.
Perhaps an easier situation to consider is the production of a secondary product
(e.g., ethanol or an antibiotic). Here we worry not so much about a mixture of subpopula-
tions, but that conditions that promote growth completely repress product formation. A
very large scale multistage system for ethanol production is currently in use. A multistage
system of CFS1R approaches PFR behavior. A PFR rnirnics the batch system, where
spaee time (the time it takes the eulture fluid to reach a specific location in the PFR) re-
places eulture time. A multistage system is much like taking the batch growth curve and
dividing it into sections, with each section being "frozen" in a corresponding stage of the
multistage system. As in the batch reactor, the culture's physiological state progresses
from one stage to the next.
The mathematical ana1ysis of the multistage system that we present here is imper-
fect. Growth in the second and subsequent stages is intrinsica1ly unbalanced growth, even
though it is steady-state growth. New cells entering the second or subsequent stage are
continuously adapting to the new eonditions in that stage. Consequently, unstructured
models are not expected to give completely accurate predictions. However, we use un-
structured models here due to their simplicity and to illustrate at least some aspects of
multistage systems.
A two-stage chemostat system is depicted in Fig. 9.3. Biomass and substrate bal-
ances on the first stage yield the following equations (ignoring endogeneous metabolism):
=7.3 gll
,.
5=~= (1)(0.065) =0.48 g/I
11m -Ilne< 0.2- 0.065
Xl = D(50 - 5)Y~/s = (0.1)(10- 0.48)0.5
Ilg 0.065
A biomass balance aroundthe concentratoryieIds
(1 + a)XI = aCXI +X2
X2 =(1 + a)XI - aCXI
= (1.7)(7.3)- (0.7)(1.5)(7.3)
=4.8 gIl
250
Then
feed is 50= 10g glucosell.The kinetic constants of the organisms are 11m = 0.2 h-l, Ks = I g
glucosell.The value of C is 1.5,and the recyc1eratio is a = 0.7. The systemis at steady
a Find the substrateconcentrationin the recyc1estream (S).
b. Find the specificgrowthra1e(Iloe~ofthe organisms.
c. Fmdthe cell (biomass)concentrationin the recyc1estream·
d. Find the cell concentrationin the centrifugeeffluent(Xv·
Solutlon Using eq. 9.9, we deterrnineIloe"
Ilnet
= [l + <XCI - C)]D= [l + (1 - 1.5)0.7](0.1)= Ilg
=0.065h-1
9.3.2. Multistage Chemostat Systems
In some fermentations, particularly for secondary metabolite production, the
product-formation steps need to be separated, since optimal conditions for each
different. Conditions such as temperature, pH, and limiting nutrients may be
each stage, resulting in different cell physiology and cellular products in m'
systems.
AD. example of a multistage system that may be beneficial is in the cultlfe
cally engineered cells. To improve genetic stability, a plasmid-carrying recombi
usually uses an inducible promoter to control production of the target protein
ter 8). ln the uninduced state, the plasmid-containing cell grOWSat nearly the s
the cell that loses the plasmid (a revertant), so the plasmid-free cell holds lil
advantage over the plasmid-containing cell. However, if the inducer is ad
mid-containing cells wi11make large quantities of the desired protein
have greatly reduced growth rates. Thus, a single-stage chemostat would fi'
for the production of the target protein because of resulting problems in gene'
A multistage system can circumvent this problem. ln the fust stage, 110 indu'
and the plasmid-containing cell can be maintained easily (usually an antibiotiC,
kill plasmid-free cells; -.s~e5napter 14 for a more complete discussion). In
Bioreaetors for Suspension and Immobilized Cultur

(9.25a
(9.26a)
(9.24t
(9.25b)
(9.26b)
9n =~= Sn -Sn-I
D --
n T
,.n
Tp.n(Xn.Sn.···) = Dn(P" - Pn-l)
8 =~ Xn-Xn_l
n Dn Tx.n(Xn,Sn)
1
T,.n =y-Tx.n(Xn,Sn) = Dn(Sn -Sn_I)
x/s
or
or
or
Datafor theproductionof a secondarymetabolitefrom a small-scalebatchreactorareshownin
Fig, 9.4, Assumethattworeactors,eachwith700-1 workingvolume.are available.Youwilluse
exactlythe same cultureconditions(medium,pH. temperature.and so on) as in the batchreac-
tor,Ifthe flowrate is 100 l/h.predictthe outletconcentrationofthe product.Comparethatto the
valuepredictedif a single 1400-1 reactorwere used, Useboth graphicalapproaches,
9 = 2.- = __1',,_-_1',,_-_1
n Dn Tp,n(Xn,SR"")
where Dn = FlVm 9n is the mean residence tÍIDein the nth stage, and Tx,n' T',n' and Tp,na11
represent rates of reaction in the nth stage.
The preceding set of equations lends itself to machine ca1culations. However,
graphica1 approaches to multistage design can a!so be used and have the advantage that
the functional form of the growth or production rate need not be known, AlI that is re-
quired is a batch growth curve. However, the transfer of the information from batch
growth curve to predictions of the multistage system stilI requires the assumption of ba!-
anced growth, Hence, the ana!ysis must be used with caution, ln at least one case (the pro-
duction of spores from Bacillus), this approach has made experimenta!ly verifiable
predictions ofthe performance vf a six-stage system,
The graphica! approaches make use of eqs. 9.24 to 9.26, Dne approach is to use a
plot of lI(dX/dt) versus X or lI(dP/dt) versus P derived from batch growth curves. This
corresponds to using eqs. 9.24b and 9.26b. The size of the required reactor is determined
by the area ofthe rectangle described with sides Xn - Xn_1 and lI(dX/dt) or Pn - Pn-I and
lI(dP/dt), The area of the rectangle is 9, and if F is known, V can be ca!culated. An a!ter-
native approach avoids some tria!-and-error solutions that are necessary with the first ap-
proach. This second approach requires plots of dXldt versus X and dP/dt versus P. The
intersection of the reaction curve with a line from the mass balance equation (e,g"
eq. 9.24a) determines the exit concentration of X or P, while the slope of the line deter-
mines D, and if F is known, V can be found. We iIIustrate the use of these approaches in
Example 9.2,
EJtaniple 9.2
(9.19)
(9.18)
where
~mS2
,_FJ+F' and ~2= K,+S2
D2- Vz
Equations 9.18 and 9.20 can be solved simultaneously for X2 and S2 by substitutin"
~ = ~mS/(K, +S2) in both equations or any other functiona1 farm that describes ~.
When a feed stream is added to the second stage, then the design equations chan.
The second feed stream may contain additiona! nutrients, inducers, hormones, or
hibitors. Biomass ba!ance for the secon<Ístage in this case is
F;Xj +F'X' - (F; +F')X2 +V2~2X2 = V2 áX2
dt
Substrate ba!ance for the second stage yields
F.S +F'S' -(F. +F')S - V2~2X2 = V dS2
I I o I 2 Y:Js 2 dt
Equations 9.22 and 9.23 need to be solved simultaneously for X2 and S2'
We can genera1ize these equations for a system with no additiona1 streantS
second or subsequent units. If we do a ba1ance around the nth stage on .biomass, su
and product, we find
At steady state when X' = O,eq. 9.21 becomes
D' F; XI
~2= 2---
Vz X2
Tx.n(Xx'Sn) = Dn(Xn - Xn_l)
~2 X2
S-- M
S2 = I D2 Yx/s
D2 = FlVz and ~2 = ~mS2
K,+S2
where
~2 =Dz(I-~)
At steady state, eq. 9.19 becomes
where XI/X2 < 1 and ~2 < D2•
The substrate ba1ance for the limiting substrate in the second stage is
FS -FS _~2X2 V = V dS2
I 2 Y~s 2 2 dt
At steady state, eq. 9.17 becomes

or
255
o
o t.2 .4
~ P,gll
Solution of Example 9.2 for two-stage system, each with e =7 h.
Modifying Batch and Continuous Reaetors
Figure9.5.
which corresponds reasonably c10sely to 7 h.
In this solution the reader should note that for the first stage, only solutions that exist
for XI greater than the value of X for which 1/(dXldt) is a minimum are practically obtainable.
Washout occurs if 6 Iis too small.
We can compare the result to a single-stage system with the same tota! volume as the
two-stage system (Fig. 9.6). Here the trial-and-error approach indicates for XI = 7.35 g/l that
7.35 gll . 1.9 h/g/l = 13.97 h = 14 h
The value of Pl that corresponds to Xl = 7.35 g/l is 0.10 g/l. Thus, the use of the two-stage
system in this case increased product concentration from 0.10 to 0.49 g/l.
An a!ternative graphica! approach that e1iminates the tria!-and-error aspect of the first
approach is shown in Fig. 9.7. Here eqs. 9.24a and 9.26a have been used. DI = 1/61 =
I
62=7 h=(Pz-JP--
d~/dt
By tria! and error, we find that at P2 = 0.49 g/l
62 = (0.49 g/l-0.08 gll)(17 h/g/l)
=6.97 h
8-
X,
7-
611
I
16-
X-
gll
4-
3-
2-
Figure 9A. Data for Example 9.2.Data
are forthe production of a secondaryP'
uet in batch culture.
61 =(XI-XO{~/dtjLl
4 8 12 16
t (h)
o
o
Bioreaetors for Suspension and Immobilized cultures .
7 h=X{~/dtl.
Since a sterile feed is to be used, Xo= O.
By tria! and error, we find on the graph that Xl = 7.2 g/I corresponds to
0.95 h/g/I or 7.2 g/I . 0.95 h/g/l = 6.84 h.
Given the accuracy with which Fig. 9.5 can be read, this is an acceptable sól1
product concentration that corresponds to Xl = 7.2 g/l is deterrnined from tM bal .
curve. As mustrated, Xl = 7.2 g/l is achieved at 9.4 h after inocuJation; at the sametu,j.
value for P1is 0.08 g/l. ..
The effect of the second stage on the process is detemuned by using eq. 9.2'
ing that again 62 =7 h. Thus,
Solution The fust step in using either graphical approach is to differentiate the data tn,
batch growth curve to yield dXldt and dP/dt. The differentiation of experimenta! data
magnify errors present in the originál data, so the values of dXldt and dP/dt must be
preted cautiously.
For the graphical approach illustrated in Fig. 9.5, we have plotted l/(dXldt)
X and lI(dP/dt) versus P, which corresponds to eqs. 9.24b and 9.26b. For 61 = 7 h (,
700 1/100 lib), we must detemune what value of Xl will satisfy
254
8
0.8
7
6
X 5
(g/Il
4
3
2

(9.28)
257
(9.29)
(9.30)
Figure 9.7. Solution to Example 92 using
a1temative graphica1 approach.
0.8
t 8
XI
6
4
X gll
2
o
O
0.15
1.0
dX/dt
gll-h
X=X'/V
MOdifying Batch and Continuous Reactors
0.10
dP/dt
gll-h
1.5
where 50 is the initial substrate concentration, Y~ is the yield coefficient, and Xo is the
initial biomass concentration. When biomass concentration reaches its maximum value
(Xm), the substrate concentration is very low, 5 « 50' and also Xo « X. That is,
Xm '" Y~o. Suppose that at Xm == Y~50' a nutrient feed is started at a tlow rate F, with the
substrate concentration 50' The total amOunt of biomass in the vessel is X' = IX, where V
is the culture volume at time t.The rate of increase in culture volume is
dV
-=F
dt
Integration of eq. 9.28 yields
V= Va+Ft
wbere Vo is the initial culture volume (I).
The biomass concentration in the vessel at any time t is
.2 .4 .6
P, g/J
O
X = Xo + Yts(So -S)
8
Xl
0.143 h-l. The intersection of the reaction curve with the straight line determi
Dl(Xj - Xo) = DlXl is the solutionto eq. 9.24a. For the second stage, we considerthe
tion phase and use eq. 9.26a. The predicted values of Xl and P2 are the same as in
approach. Note thal the dP/dt-versus-P curve is displaced in time from the dXldt-
Xcurve. Consequently,we use the dXldt plot before using the dP/dt plot.
Figure 9.6. Solution ofExarnple 9,2 with a single stage, where e = 14h,
2 4 6
X, O/I
ln fed-batch culture, nutrients are continuously or semicontinuously fed, while efll
removed discontinuously (Fig. 9.8). Such a system is caIled a repeatedfed-batch
Fed-batch culture is usuaIly used to overcome substrate inhibition or catabolite
by intermittent feeding of the substrate. If the substrate is inhibitory, intermitten
of the substrate improves the productivity of the fermentation by maintaining the
concentration low. Fed-batch operation is also called the semicontinuous
variable-volume continuous culture. Consider a batch culture where the concenl
biomass at a certain time is given by
9.3.3. Fed-batch Operation
256
o
O
8-
X.
7-
6- X_
"
O/I
4-
3-
2-

259
(9.43)
(9.42)
(9.41)
(9.39)
(9.40)
(9.36)
(9.35)
(9.38)
(9.37)
FSo = I! •••X'
Y~s
X' = X~+FYftsSot
Ps YPISSO
" ( Ft)
P = Po+qpXm Vo+2" t
Equation 9.31 at quasi-steady state with S ""Oyields
FP"" YPlsSoF
When the specific rate of product formation qp is constant,
dP' ,
-=qpX
dt
where P' is the tota! amount of product in culture.
Substituting X' = (Vo+ Ft)Xm into eq. 9.41 yields
dP'
-=qpXm(Vo+Ft)
dt
or the potential product output is
dX' ("dV) M
- = Xm - = XmF = FYxIsSo
dt dt
Integration of eq. 9.37 from t = O to t with the initial amount of biomass in the reactor
being Xó yields
That is, the total amount of cell in the culture increases linearly with time (which is exper-
imentally observed) in a fed-bateh cultore. Dilution rate and therefore fJ.n •• decrease with
time in a fed-bateh culture. Since !J.ne, = D at quasi-steady state, the growth rate is con-
trolled by the dilution rate. The use of unstructured models is an approximation, since fJ.n ••
is a function of time.
Produet profiles in a fed-batch culture can be obtained by using the definitions of
YPIS or qp. When the product yield coefficient YPIS is constant, at quasi-steady state with
S «So
The balance on the rate-limiting substrate without maintenance energy is
dS' = FSo _ IlnerX'
dt y:~
where S' is the tota! amount of the rate-Iimiting substrate in the cultore and Sois the con-
centration of substrate in the feed stream.
At quasi-steady state, X' = VXm and essentially all the substrate is consumed, so no
significant level of substrate can accumulate. Therefore,
Fill
Start
!J.net =D
Ss Kp
11m -D
S
I1net = 11mK+S
..
Bioreactors for Suspension and Immobilized Cultures'
F,So
llJ
SE
258
then
If maintenance energy can be neglected,
Figure 9.8.
V.' X. S, PI Harvest culture.
Tbe rate of change in biomass concentration is
dX V(dX'/dt)-X'(dV/dt)
dí V2
Since dX'/dt= 1J.ne,X', dV/dt= F, and FIV=D, eq. 9.31 becomes
dX
dí = (I1net - D)X
When the substrate is totally consumed, S ""O and X = Xm = ylfJsSo. Furthermore•
nearly all the substrate in a unit volume is consumed. then dXldt = O.This is an e
of a quasi-steady state. A fed-batch system operates at quasi-steady state when IIi
consumption rate is nearly equal to nutrient feed rate. Since dXldt = O at quasi·
state, then

261
F= dV =200 m1/h
dt
Ilm = 0.3 h-I
Y~s =0.5 gdw cellsl g glucose
v = 1000 ml
So= 100 g glucose/l
Ks = 0.1g g1ucose/l
X~ =30 g
d. P=Po Vv +q X (Vv Dt)
V P m V'+T t
= 0+(0.2)(50) (600 + (0.2)(2»)
1000 --- (2)
=16 g/l 2
c. X' = X~+ FY::sSot
=30+(0.2)(0.5)(100)(2)=50 g
a Find Vo(the initial volumeof the culture).
b. Deterrnine the concentration of growth-Iirniting substrate in the vessel at quasi-steady
state.
c. Detennine the concentration and tota! amount of biomass in the vessel at t = 2 h (at
quasi-steadystate).
d. If qp = 0.2 g product/gcells, Po = O,deterrninethe concentrationof product in the vessel at
t=2h.
Solution
a. V= Vo+Ft
Vo = 1000 - 200(2) = 600 ml
b. D= FIV= 0.2 h-1
S KsD (0.1)(0.2) O2 I II
~--- = - . ggucose
Ilm - D 0.3-0.2
tions such as lactic acid and other plant cell and mamma!ian cell fermentations. where the
rate of product formation is maxima! at low nutrient eoncentrations.
Fed-batch culture is important for E. coli fermentations to make proteins from reeom-
binant DNA techDology. To make a high concentration of product, it is desirable to grow the
culture to very high cell density before inducing production of the target protein. If E. coli
has an unlimited supply of glucose it will grow at a maxima! rate, but produce organic acids
(e.g., acetic acid) as by-products. The accumulation of these by-products inhibits growth.
If glucose is fed at a rate that substains the growth rate at slightly less than maxima!, E. coli
uses the glucose more efficiently, making less by-product. Very high cell densities (50 to 100
gll) can be achieved. Fed-batch culture may benefit from active process contro!. For exam-
ple, the feed rate of glucose could be controlled by measuring glucose concentration in the
medium or the CO2 evolution rate using a feedback controller.
Example9.3
In a fed-batch culture operating with intennittent addition of glucose solution. values of the
followingparametersare givenat time t = 2 h. whenthe systemis at quasi-steadystate.
FiguN 9.9. (a)Variation
ofcul!
(V) • specific
growth
rate(I.l). ceIl
substrate
(S) concentration
withIÍ
quasi-steady
state.(b)Variation
of.,
(P) concentration
withtimeatq .
statein a singlecycleofafed-b
Time
(b)
p'Ypx
p(qpCOllIIonII
L.
Time
Bioreaetors for Suspension and Immobilized CulturElS
x
So
260
Substitution of eq. 9.46 into eq. 9.45 yields
Pw = rPo+ q;~m (1-r2)
w
An example of fed-bateh eulture is its use in some antibiotie fennentations, w.
glucose solution is intermittently added to the fennentation broth due to the repressi'
pathways for the production of seeondary metabolites eaused by high initia! glueose
eentrations. The fed-batch method ean be applied to other secondary metabolite fl
In tenns ofproduet eoneentration, eq. 9.43 ean be written as
P=Po ~ +qpXm(~ +~}
Figure 9.9 depiets the variation of V, !J. (= D), X, 5, and P with time at quasi-steady
state in a single eyc1e of a fed-bateh eulture.
In some fed-bateh operations, part of the eulture volume is removed at eertain inter·
va!s, sinee the reaetor volume is limited. This operation is ea!led the repeated fed-batch
culture. The eulture volume and dilution rate (= llneJ undergo eyeliea! variations in this
operation.
If the eycle time tw is eonstant and the system is a!ways at quasi-steady state, then
the produet coneentration at the end of eaeh eycle is given by
Pw=rPo +qpxm(r + D~w }w
where Dw = FlVw' Vw is the eulture volume at the end of each cycle, Vo is the residua! c
ture volume after remova!, y is the fraetion of eulture volume remaining at eaeh cY'
(= VdVw)' and tw is the eycle time.
The eycle time is defined as
Vw-Vo Vw-yVw l-y
t == ~ =--,,--,---",--
--
w F F Dw

Product
263
Immobilized Ce" Systems
.9,4
1. Immobilization provides high cell concentrations.
2. Immobilization provides cell reuse and eliminates the costly processes of cell recov-
ery and cell recycle.
3. Immobilization eliminates cell washout problems at high dilution rates.
4. The combination of high cell concentrations and high flow rates (no washout re-
strictions) allows high volumetric productivities.
5. Immobi1ization may also provide favorable microenvironmental conditions (Le.,
cell-cell contact, nutrient-product gradients, pH gradients) for cells, resulting in
better performance ofthe biocatalysts (e.g., higher product yields and rates).
6. In some cases, irnmobilization improves genetic stability.
7. For some cells, protection against shear damage is important.
9.4.2. Aetive Immobilization of Cells
The major limitation on immobilization is that the product of interest should be ex-
creted by the cells. A further complication is that immobilization often leads to systems
for which diffusionallimitations are important. In such cases the control of microenviron-
mental conditions is difficult, owing to the resulting heterogeneity in the system. With liv-
ing cel!s, growth and gas evolution present significant problems in some systems and can
lead to significant mechanical disruption of the immobilizing matrix.
In Chapter 3 we discussed enzyme immobilization. Figure 3.16 provides a useful
summary of immobilization strategies. Many of the ideas in enzyme immobilization have
a direct counterpart in whole cells. However, the maintenance of a living cell in such a
system is more complex than maintaining enzymatic activity. The primary advantage of
irnmobilized cells over immobilized enzymes is that immobi1ized cells can perform multi-
step, cofactor-requiring, biosynthetic reactions that are not practical using purified en-
zyme preparations.
9.4.1. Introdu~ion
ImmobiIization of cells as biocatalysts is almost as common as enzyme immobilization.
Immobilization is the restriction of cell mobility within a defined space. Immobilized cell
cultures have the following potential advantages over suspension cultures.
Active immobilization is entrapment or binding of cells by physical or chemical forces.
The two major methods of active immobilization are entrapment and binding.
Physical entrapment within porous matrices is the most widely used method of celI
. inunobilization. Various matrices can be used for the immobilization of cells. Among
.'these are porous polymers (agar, alginate, K-carrageenan, polyacrylarnide, chitosan,
'gelatin, collagen), porous metal screens, polyurethane, silica gel, polystyrene, and cellu-
'lOsetriacetate.
9.4. IMMOBILlZED CELL SYSTEMS
Supernotert ,
Woste
a
Oead Cell~
Live CeUReturn
Bioreactors for Suspension and Immobilized CulturéS
Medium
Reservoir Bioreoctor
Fig. 9.10. Schematic of a perfusion system with externa! centrifugation and returll
cells. Internal retention of cel1s is a!so possible. ReturD of spent medium is optional.
262
9.3.4. Perfusion Systems
An alternative to fed-batch culture is a perfusion system. Such systems are used most
often with animal cell cultures (see Chapter 12). The basic characteristic is constant
medium {low, cell retention, and in some cases selective removal of dead cells. High cel!
density can be achieved. Cell retention is usually achieved by membranes or screens or by
a centrifuge capable of se1ective cell removal. When a membrane is used, the system has
characteristics of an immobilized cell system (see Section 9.4) except the cells are usual!y
maintained in suspension and rni.J(ed.With a selective removalJrecycle the system ap-
proaches the cell recycle reactor discussed earlier in this chapter. Figure 9.10 depicts one
type of perfusion system.
The potential advantages of a perfusion system is the potential removal of cell
bris and inhibitory by-products, removal of enzymes released by dead cells that may 00'
stroy product, shorter exposure time of product to potentially harsh production conditio
(compared to batch or fed-batch operation), high per-unit volumetric productivity (due
high cell density and metabolism), and a rather constant environment.
The primary disadvantage is that a large amount of medium is typically used
the nutrients in the medium are less completely utilized than in batch or fed-batch
tems. High medium usage is expensive, owing not only to the high cost of raw mate
but also to the costs to prepare and sterilize the medium. Additionally, costs for
treatrnent increase. Typically the bioprocess engineer must consider the trade-off of .
proved product quality and reactor productivity with the extra costs associated wil
more complex reactor system (membranes, pumps, centrifuga! separator, etc.)
creased medium usage. The best choice depends on the specific situation.

265
Immobilized Ce/l Systems
Encapsulation is another method of cell entrapment. Microcapsules are hollow,
spherical particles bound by semipermeable membranes. Cells are entrapped within the
hollow capsule volume. The transport of nutrients and products in and out of the capsule
takes place through the capsule membrane. Microcapsules have certain advantages over
gel beads. More cells can be packed per unit volume of support material into capsules,
and intraparticle diffusion limitations are less severe in capsules due to the presence of
liquid cell suspension in the intracapsule space. Various polymers can be used as capsule
membranes. Among these are nylon, collodion, polystyrene, acrylate, polylysine-alginate
hydrogel, cellulose acetate-ethyl cellulose, and polyester membranes. Different mem-
branes (composition and MW cutofi) may need to be used for different applications in
order to retain some high-MW products inside capsules and provide passage to low-MW
nutrients and products.
Another form of entrapment is the use of macroscopic membrane-based reactors.
.The simplest of these is the hollow-fiber reactor. This device is a mass-transfer analog of
the shell-and-tube heat exchanger in which the tubes are made of semipermeable mem-
branes. Typically, cells are inoculated on the shell side and are allowed to grow in place.
The nutrient solution is pumped through the insides of the tubes. Nutrients diffuse through
the membrane and are utilized by the cells, and metabolic products diffuse back into the
flowing nutrient stream. Owing to diffusional limitations, the unmodified hollow-fiber
unit does not perform well with living cells. Modifications involving multiple membrane
types (for example, for gas exchange or extractive product removal) or changes to pro-
mote convective flux within the cell layer have been proposed. Several commercial reac-
tors for animal cell cultivation use membrane entrapment.
ln addition to entrapment or encapsulation, cells can be bound directly to a support.
Immobilization of cells on the surfaces of support materials can be achieved by physical
adsorption or covalent binding.
Adsorption of cells on inert support surfaces has been widely used for cell immobi-
lization. The major advantage of immobilization by adsorption is direct contact between
nutrient and support materials. High cellloadings can be obtained using microporous sup-
port materials. However, porous support materials may cause intraparticle pore diffusion
limitations at high cell densities, as is also the case with polymer-entrapped cell systems.
Also, the control of microenvironmenta! conditions is a problem with porous support ma-
terials. A ratio of pore to cell diameter of 4 to 5 is recommended for the immobilization of
cells onto the inner surface of porous support particles. At small pore sizes, accessibility
of the nutrient into inner surfaces of pores may be the limiting factor, whereas at large
pore sizes the specific surface area may be the limiting factor. Therefore, there may be an
optimal pore size, resulting in the maximum rate of bioconversion.
Adsorption capacity and strength of binding are the two major factors that affect the
selection of a suitable support material. Adsorption capacity varies between 2 mg/g
(porous silica) and 250 mg/g (wood chips). Porous glass carriers provide adsorption ca-
pacities (lOS to 109 cells/g) that are less than or comparable to those of gel-entrapped cell
concentrations (109 to 1011 cellslml). The binding forces between the cell and support sur-
faces may vary, depending on the surface properties of the support material and the type
of cells. Electrostatic forces are dominant when positively charged support surfaces (ion-
exchange resins, gelatin) are used. Cells also adhere on negatively charged surfaces by co-
valent binding or H bonding. The adsorption of cells on neutral polymer support surfaces
264
Polymer beads should be porous enough to allow the transport of substrates and
products in and out of the bead. They are usually formed in the presence of cells and can
be prepared by one of the following methods:
1. Gelation oj polymers: Gelatin and agar beads may be prepared by mixing the
liquid form of these polymers with cell suspensions and using a template to form beads.
Reduction of temperature in the templates causes solidification of the polymers with the
cells entrapped. Gel beads are usually soft and mechanically fragile. However, we can use
a hard core (glass, plastic) and a soft gelatin shell with entrapped cells to overcome some
mechanical problems associated with polymer beads. Because of diffusionallimitations,
the inner core of such beads is ofted not active, so this approach does not necessarily de-
crease the amount of product made per bead.
2. Precipitation oj polymers: Cells are dispersed in a polymer solution, and by
changing the pH or the solvent, the polymer can be precipitated. The starting solution of
the polymer has to be prepared with an organic solvent or a water-solvent mixture.
Ethanol and acetone are examples of water-miscible solvents. polymers used for this pur- .
pose are polystyrene, cellulose triacetate, and collagen. The direct contact of cells with'
solvents may cause inactivation and even the death of cells.
3. lon-exchange gelation: lon-exchange gelation tak.es place when a water-solub
polyelectrolyte is mixed with a salt solution. Solidification occurs when the polyeI
trolyte reacts with the salt solution to form a solid gel. The most popular example of
kind of gelation is the formation of Ca-alginate gel by mixing Na-alginate solution wi'
CaClz solution. Some other polymers obtained by ion-exchange gelation are Al-algin
Ca/Al carboxymethyl cellulose, Mg pectinate, lC-carrageenan, and chitosan pol
phosphate. Alginate and lC-carrageenan are the most widely used polymers for ce'
irnmobilization purposes. lonic gels can be further stabilized by covalent cross-linking.
4. Polycondensation: Epoxy resins are prepared by polycondensation and can
used for cell immobilization. polycondensation produces covalent networks with ..
chemical and mechanica! stability. Usually, liquid precursors are cured with a mulu
tiona! component. Functional groupS usua!ly are hydroxy, arnino, epoxy, and isoc:
groups. Some examples of polymer networks obtained by polycondensation are e
polyurethane, silica gel, gelatin-glutaraldehyde, albumin-glutaraldehyde, and collag'
glutaraldehyde. Severe reaction conditions (high temperature, low or high pH valuesJ
toxic functiona! groupS may adversely affect the activity of cells.
5. Polymerization: Polymeric networks can be prepared by eross-linking C' .
mers of a vinyl group containing monomers. polyacrylarnide beads are the most
used polymer beads, prepared by copolymerization of acrylarnide and bisacrylarni'
eral different monomers can be used for polymer formation; acrylamide, methacryl:
and 2_hydroxyethyl methacrylate are the most widely used. Cross-linking is usua1l:
ated by copolymerization with a divinyl compound, such as methylenebis-acrylarni'
Immobilization by polymerization is a simple method. The polymerizing sol
mixed with the cell suspension, and polymerization takes place to form a polymeric
which is pressed through a sieve plate to obtain regular-shaped particles. Suspen:
emulsion polymerization can also be used to form polymeric beads for cell entraptn'

9.4.3. Passive Immobilization: Biological Films
Conversion
Glucose to lactic acid
Glucose to acetone, butanol
Streptomycin
Hormones
267
Steroid glycoalkaloids formation
Glucose to ethanol
Glucose to 2,3-butanediol
Fumaric acid to aspartic acid
Cellulose production
Glucose to ethanol
Glucose to glucomc acid
Anthraquinone formation
Phenol degradation
Conversion of testosterone
Pemcillin G to G-APA
lsocitrate dehydrogenase aetivity
Menthyl succinate to menthol
Conversion
Support surface
Gelatin (adsorption)
Ion-exchange resins
Sephadex (adsorption)
DEAE-sephadexlcytodex (adsorption)
Ti(lV) oxide (covaient binding)
Agarose-cartJodiimide (covalent
binding)
Polyphenylene oxide-glutaraidehyde
(covalent binding)
K-Carrageenan or polyacrylamide
K-Carrageenan
K-Carrageenan
K-Carrageenan
Ca-alginate
Ca-alginate
Ca-alginate
Ca-alginate
Polyurethane
Polyurethane
Polyurethane
Polyurethane
Support matrix
Immobilized Cell Systems
Cells
S. cerevisiae
E. aerogenes
E. coli
Trichoderma reesei
Z. mobilis
Acetobacter sp.
Morinda citrifolia
Candida tropicalis
Nocardia rhodocrous
E. coli
Catharantus roseus
Rhodotorula minuta
Cells
TABLE 9.2 Examples of Celllmmobilization by Surface Attachment
Lactobacillus sp.
Clostridium acetobutylicum
Streptomyces
Anima! cells
E.coli
B. subtillis
Solanum aviculare
ln mixed-culture microbia! films, the presence of some polymer-producing organ-
isms facilitates biofilm formation and enhances the stability of the biofiIms. Microenvi-
ronmental conditions inside a thick biofiIm vary with position and affect the physiology
of the cells.
ln a stagnant biologica! film, nutrients diffuse into the biofiIm and products diffuse
out into liquid nutrient medium. Nutrient and product profiIes within the biofilm are im-
portant factors affecting cellular physiology and metabolism. A schematic of a biofilm is
depicted in Fig. 9.11. Biofilm cultures have a!most the same advantages as those of the
immobiIized celI systems over suspension cultures, as Iisted in the previous section.
The thickness of a biofiIm is an important factor affecting the performance ofthe bi-
otic phase. Thin biofilms wiII have low rates of conversion due to low biomass concentra-
tion, and thick biofilms may experience diffusionally limited growth, which may or may
not be beneficial depending on the cellular system and objectives. Nutrient-depleted re-
gions may also develop within the biofilm for thick biofilms. In many cases, an optima!
biofiIm thickness resulting in the maximum rate of bioconversion exists and can be deter-
mined. In some cases, growth under diffusion limitations may result in higher yields of
products as a result of changes in celI physiology and cell-eelI interactions. In this case,
TABLE 9.1 Examples of Celllmmobilization by Entrapment Using Different
Support Materials
Sec. 9.4
Bioreaetors for Suspension and Immobilized Cultures
266
Biological films are the multilayer growth of cells on solid support surfaces. The sup
material can be inert or biologica!ly active. Biofilm formation is common in natural
industrial fermentation systems, such as biological waste-water treatment and mold
mentations. The interaction among cells and the binding forces between the celI and s
port material may be very complicated.
may be mediated by chemical bonding, such as cova!ent bonding, H bonds, or van der
Waals forces. Some specific chelating agents may be used to develop stronger celI-surface
interactions. Among the support materials used for celI adsorption are porous glass,
porous siIica, alumina, ceramics, gelatin, chitosan, activated carbon, wood chips,
polypropylene ion-exchange resins (DEAE-Sephadex, CMC-), and Sepharose.
Adsorption is a simple, inexpensive method of celI immobiIization. However, lim-
ited celI loadings and rather weak binding forces reduce the attractiveness of this method.
Hydrodynamic shear around adsorbed cells should be very mild to avoid the removal of
cells from support surfaces.
Cova!ent binding is the most widely used method for enzyme immobilization. How·
ever, it is not as widely used for celI immobiIization. Functional groups on celI and sup-
port materia! surfaces are not usually suitable for cova!ent binding. Binding surfaces need
to be specialIy treated with coupIing agents (e.g., glutara!dehyde or carbodiimide) or reac-
tive groups for cova!ent binding. These reactive groups may be toxic to cells. A number of
inorganic carriers (metal oxides such as titanium and zirconium oxide) have been devel·
oped that provide satisfactory functional groups for cova!ent binding.
Cova!ent binding forces are stronger than adsorption forces, resulting in more stablé
binding. However, with growing celIs, large numbers of cell progeny must be lost. Sup-
port materials with desired functional groups are rather limited. Among the support mate·
ria!s used for covalent binding are CMC plus carbodiimide; carriers with aldehyde, amine,
epoxy, or ha!ocarbonyl groups; Zr(IY) oxide; Ti(IY) oxide; and cellulose plus cyanuric
chloride. Support materia!s with -OH groups are treated with CNBr, materia!s with
-NH2 are treated with glutaraldehyde, and supports with COOH groups are treated with
carbodiimide for covalent binding with protein groups on cell surfaces.
The direct cross-linking of cells by glutara!dehyde to form an insoluble aggregate .
more like cell entrapment than binding. However, some cells may be cross-linked
adsorption onto support surfaces. Cross-Iinking by glutaraldehyde may adversely affect,
the cell's metabolic activity and may also cause severe diffusion limitations. Physica!
cross-linking may a!so be provided by using polyelectrolytes, polymers such as chitos
and salts [CaCI2,Al(OH)3,FeCI3]. Direct cross-linking is not widely used because of
aforementioned disadvantages.
A good support materia! should be rigid and chernically inert, should bind ce
firmly, and should have high loading capacity. In the case of gel entrapment, gels shouJ,
be porous enough and particIe size should be smalI enough to avoid intraparticIe diffusio!
limitations.
Some examples of cell immobilization by entrapment and by surface attachm
(binding) are summarized in Tables 9.1 and 9.2, respectively.

269
(9.48)
(9.49)
D d2; =_1_ JlmS X
e dy YX1S Ks +S
Da = maximum rate of bioconversion = rmax
maximum rate of diffusion (D/o)So
Sec. 9.4
where rmax is the maximum rate of bioconversion (mg Sil h), De is the effective diffusivity
of the rate-limiting substrate, o is the thickness of diffusion path (or liquid film), and Sois
the bulk substrate concentration in liquid phase. When the film-theory model applies,
D/o is the mass transfer coefficient (Le., kL = D/o).
If Da » I, the rate of bioconversion is diffusion limited; for Da « I, the rate is
limited by the rate of bioconversion; and for Da '" I, the diffusion and bioreaction rates
are comparable. It is desirable to keep Da < I to eliminate diffusion limitations when the
productivity of a cell population does not improve upon immobi1ization due to cell-eell
contact and nutrient gradients.
Diffusiona!limitations may be extemal (that is, between fluid and support surface in
adsorption and cova!ent binding), intrapartic1e (Le., inside partic1es in entrapment, encap-
su1ation, or irnmobi1ization in porous partic1es), or both. If the extema! mass transfer is
limiting, an increase in liquid-phase turbulence should result in an increase in the reaction
rate. In case of intrapartic1e mass-transfer limitations, a reduction in partic1e size or an in-
crease in the porous void fraction of the support materia! should result in an increase in
the rate of the bioreaction.
In Chapter 3 we discussed in reasonable detail a mathematica! model of the interac-
tion of diffusion and reaction for surface immobi1ized or entrapped biocatalysts. These
model s apply directly to immobi1ized cells when the kinetics of bioconversion are de-
scribed by a Michaelis-Menten type of kinetic expression. Thus, the reader may wish to
consult Chapter 3 again.
Another interesting case is to consider biofilms where we allow cell growth. Models
for immobi1ized enzymes have no terms for biocata1yst replication, so this case presents a
new problem.
The thickness of a biofilm or the size of microbia! floc increases with time during
the growth phase. A microbial floc is an aggregation of many cells, and in some processes
these aggregates can be more than 1 mm in diameter. However, since the rate of increase
in biofi1m thickness is much slower than the rate of substrate uptake, the system can be
assumed to be at quasi-steady state for relatively short periods. The simplest case is to
assume that the system is at quasi-steady state and a!1 the cells inside the biofilm are
in the same physiological state. In this situation we write a steady-state substrate bal-
ance within the biofi1m by using average kinetic constants for the biotic phase (living
cells).
A differential materia! ba!ance for the rate-limiting substrate within the biofilm (see
Fig. 9.11) yields at steady state
where De is the effective diffusivity (cm2/S) and YXIS is the growth yield coefficient
(g cells/g substrate).
Figure 9.12. Dissolved-oxygen pro
and oxygen gradients in a microbial :.,
bathed in flowing medium: -A-A- ox
profi1e for 20 ppm nutrient broth, 27.5'
_ -{)xygen gradient for this profi1e; -,
oxygen profi1e for 500 ppm nutrient bi
260C; oxygen gradient for this prolíl
(With pennission, from H. R. Bungay .
others, Bioteehno/. Bioeng. 1I:765, 19'
John Wl1ey & Sons, Inc., New Yor<.)
Figure 9.11 Schematic representation
biofilm.
0'07
e
0'02 J..
0·03l
li
0'04
O'OS !
0'06S
0'01
0'00
100 50 o 50
Ibofttum
Dlstance (MlD)
268
1
o
improvement in reaction stoichiometry (e.g., high yield) may.overcome the reduction in
reaction rate, and it may be more beneficia! to operate the system under diffusion limita-
tions. Usua1ly, the most sparingly soluble nutrient, such as dissolved oxygen, is the:
rate-limiting nutrient within the biofilm. A typica! variation of dissolved oxygen concen-.
tration within the biofilm is depicted in Fig. 9.12.
9.4.4. Diffusional Limitations in Immobilized Cell Systems
Immobi1ization of cells may cause extra diffusiona!limitations as compared to suspensi,
cultures. The presence and significance of diffusiona! limitations depend on the relat"
rates of bioconversion and diffusion, which can be described by the Damkohler numl
(Da) (see eq. 3.52 a!so).
OS
"a
i4
i3
8
~ 2
sr
1
7
16
...••

(9.55)
(9.56)
(9.57)
271
(9.58)
Figure 9.13. Effectiveness factor for a fiat
biofilm as a function of ~, the dimensionless
initial substrate concentration, and ljl, the
Thiele modulus. (With pennission, redrawn
from B. Atkinson, Biochemica/ Reactors,
Pion Ltd., London, 1974, p. 81.)
w=.:&
Ks
1000
r
r==-,
R
100
d2S 2 dS <jl2S
-+--=-
dr2 r dr I+S/W
S=~
s.'
o
Immobilized CeIJSystems
(<jl< 1) to elirninate diffusion limitations. As the biofilm grows (slowly), the value of <jl
will gradually increase. If shear forces cause a portion of the film to detach, then <jlwill
decrease abruptly.
The effectiveness factor (11)can be calculated as
11= 1 - tanh<jl(~ _ I), for 00:;;;
1
<jl tanhOl
1 tanh<jl( 1 J
11=00 - -<jl- tanh<jl - 1, for 00;:::1
where 00is the modified Thiele modulus and is given by
00= <jl(SoIKs) [So _ In(I+~Jr/2
~(1+
~J Ks Ks J
Some cells such as molds (A. niger) form pellets in a fermentation broth, and sub-
strates need to diffuse inside pellets to be available for rnicrobial consumption. Cells may
form biofilms on spherical support particles, as depicted in Fig. 9.14. Sirnilar equations
need to be solved in spherical geometry in this case to determine the substrate profile
within the floc and the substrate consumption rate. The dimensionless substrate transport
equation within the microbial floc is
where
Sec. 9.4
~=~
Ks
aty = L
aty = O
- y
y=r;'
IlmX _ L ~ rm
YXISDeKs DeKs
S=SQi
dS =0
cly
- S
s=-
S'o
<jl=L
The boundary conditions are
where
270
and
where L is the thickness of biofilm.
If it is also assumed that the liquid nutrient phase is vigorously agitated and the liq-
uid film resistance is negligible, then So '" Soi' By defining a maximum rate of substrate
utilization as rm = Il",x/YXIS (g subs/cm' h), we rewrite eq. 9.49 as
"Dd2S=~
e dl Ks+S
In dimensionless form, eq. 9.50 can be written as
d2S <jl2S
di = 1+~S
Equation 9.51 can be solved numerically. An analytical solution can be derived for
lirniting cases of zero or first-order reaction rates.
The maximum rate of substrate flux in the absence of diffusion lirnitations is
by the following equation:
N A = -A D dS I = rmSO (LA)
s s sed y=O K + S s
y s o
N --D dSl - ( rmSO JL
s- edy y=o-11 Ks+So
where 11is the effectiveness factor, defined as the ratio of the rate of substrate con:
tion in the presence of diffusion lirnitation to the rate of substrate consumption in
sence of diffusion limitation. In the absence of diffusion limitations, 11== 1, and
presence of diffusion limitations, 11< 1. The effectiveness factor is a function of <jl
Figure 9.13 is a plot of 11versus ~ for various values of <jl.The <jlvalue should
where As is a surface area of biofilm available for substrate flux, Ns is the substrate
and L is the thickness of the biofilm.
In the presence of diffusion lirnitation, the rate of substrate consumption or f1,
expressed in terms of the effectiveness factor.

27~
(9.65)
(9.64)
(9.67)
(9.68)
r = /lmX
s ---=1:
Yx/s m
S=So --.!Í!L(K _(2)
6D.
,.9.4
1] = 1-(1- 6D.So J312
rmR2
9.4.5. Bioreactor Considerations ln Immobilized
Cel! Systems
or
Various reactor configurations can be used for immobilized cell systems. Since tbe sup-
port matrices used for cel! immobilization are often mechanical!y fragile, bioreactors witb
low hydrodynamic shear, such as packed-column, fluidized-bed, or airlift reactors, are
preferred. Mechanically agitated ferrnenters can be used for some immobilized-cell sys-
tems if tbe support matrix is strong and durable. Any of these reactors can usual!y be op-
erated in a perfusion mode by passing nutrient solution through a column of immobilized
cells. Schemafic diagrams of immobilized-cell packed-column and fluidized-bed reactors
are depicted in Fig. 9.15. Tbese reactors can be operated in batch or continuous mode.
Consider tbe reactors shown in Fig. 9.15. When tbe fluid recirculation rate is high,
lesystem approaches CFSTR behavior. One commercial fluidized-bed, immobilized-
hnal-cel! bioreactor system requires high recirculation to maintain uniforrn conditions
tbe reactor. Tbe models we have discussed so far can be applied to such systems. Tbe
Substrate concentration may be zero at a certain radial distance from tbe center of
tbe floc according to eq. 9.65. This distance is called tbe critical radius (rcr) and is deter-
mined by setting S = O at rá in eq. 9.65.
(rcr)2 =1- 6D.So (9.66)
R rmR2
When rcr > o-tbat is, R > (6D .501rm)'a_tben tbe concentration of tbe limiting sub-
strate is zero for O < r < rcr In this case, tbe limiting substrate is consumed only in tbe
outer shell of the floc, and tbe effectiveness factor is given by
4 3 3
rm-1t(R -rcr) (r )3
1]= 3 =1- ..sr.
~1tR3.r R
3 m
Tbe solution to eq. 9.58 in this case is
Tbe rate of bioreaction can be approximated to zero order at values of S » Ks. Be-
cause Ks is often very small, the zero-order limit useful!y describes many systems of prac-
tical interest.
Figure 9.14. (a) Microbial film on inert
spherical support particle. (b) Spherical
microbial floc.
50
$= Vp ~rmIK=
Ap D.
I!mX =Rt rm=
Yx1sD.Ks D.Ks
"
<>=R
Bioreactors for Suspension and Immobilized CulturéS'
50
b)
272
where
Tbe boundary conditions are
8=1, atr=1
d8 O - O
_= atr=
dr '
For nonspherical partic1es, a characteristic lengtb is defined as
L= Vp
Ap
where Vp and Ap are tbe volume and surface area of rnicrobial pellet.
The rate of substrate consumption by a single rnicrobial floc is
NsAp=-ApD dS =1]~VP
'drr=R Ks+So
The effectiveness factor (1]) is a function of <>
and ~. Variation of 1] witb <>
and ~ is
to that of Fig. 9.13. However, 1] values for spherical geometry are slightly 10W'
those of rectangu1ar geometry for intermediate values of <ji(1 < <>< 10). An ana1yti'
lution to eq. 9.58 is possible for first- and zero-order reaction kinetics.
The reaction rate can be approximated to fust order at low substrate conceh!
"S 1:
r =.J:!!l=-X =...!!!-S
s Yx/sKs Ks
where rm = (I!JY XlS)X. The effectiveness factor in this case is given by
1[ 1 11
1]=$L~-3<>j
and
o)

_FdSo =11 rmSo LaA
dz K, +So
.•.•..
(9.72)
ln ~ = _ 'f1rmLaA
SOj ~H s
(100- 2) = (0.2)(25) (1r/4)(10/ 8
0.49 400
8=49 dm=4.9 m
Immobilized Cel! Systems
c. p,= Yp/S (So;- So) = 0.49(98) = 48 gIl.
; 9.4
For low substrate concentrations in the feed, the rate of substrate consumption is
fust order and eq. 9.71 has the following form:
Example9A
Glueose is converted to ethanol by immobilized S, cerevisiae eens entrapped in Ca-alginate
beads in a packed column. The specifie rate of ethanol production is qp = 0.2 g ethanol/
g cen, h, and the average dry-weight cen coneentration in the boo is X = 25 gIl bed, Assume
that growth is negligible (Le., almost all glucose is converted to ethanol) and the bead size is
sufficiently small that TI :: l. The feed flow rate is F = 400 l/h, and glueose concentration in
the feOOis SQj= 100 g glueose/L The diameter of the eolumn is 1 m, and the produet yield co-
efficient is YPIS '" 0.49 g ethanol/g glucose.
a. Write a material balance on the glucose concentration over a differential height of the col-
umn and integrate it to determine S = S(z) at steady state.
b. Determine the column height for 98% glucose eonversion at the exit of the column.
c. Determine the ethanol concentration in the effluent.
Solution
a. A material balance on the glucose concentration over a differential height of the column
yields
Substrate concentration drops exponentially with the height of tbe column in this case,
and a plot of ln So versus H results in a straight line. Equation 9.71 or 9.72 can be used as
tbe design equation for immobilized-biofilm column reactors to determine the height of
the column for a desired level of substrate conversion.
so,.-s _ qpX A
0---8
YPIS F
s - H
-Ff. .dSo = qpX Ai dz
5", Yp1S o
lntegration yields
-FdSo=qpXdV=qpx Adz
YplS YP1S
This equation differs from the form of eq. 9.72 because So;is high and the reaction rate is
effectively zero order.
b. So= 0,02(100) = 2 g glucoselL Substituting the given values into the above equation yields
Recycle
Chomber
r- Feed from
Reservoir
I
Feed from
Reservoir
"
Recycle
Chomber
Pump
t Pump
.
:
.
Alternollve
Air lnlel
Figure 9.15. Sehematies of a paeked-bed and a fluidized-bed biofilm or immobilized-
cen bioreaetors are shown, In batch operation. only tbe sucams witb soUd Unes exist. In
continuous operauon. tbe streams shown by dashed Unes are added, For the fluidized bed,
fluidizauon can be aceompUshed by Uquid reCÍfcnlation only or a ffiÍXture of liquid and
gas flows,
274
Integration of eq. 9.70 yields
K ln~+(So' -So)= Tl'mLaA H
'So I F
where SOi is the inlet bulk substrate concentration, L is tbe biofilm thicknesS or
teristic 1engtb of tbe support particle (L = Vp/Ap), and H is tbe total height of
bed.
otber extreme involves some waste-treatment systems where tbe rate of fluid recircul:
is low or even zero. ln tbe latter case, Ihe system cannot be modeled as a CFSTR but
be treated as a PFR. To analyze such a system, eonsider a materlal balance on tbe
limiting substrate over a differential element:
-F tiSo =NsaA dz
where So is tbe bulk liquid-phase substrate concentration (mg S/cm3) and is a funeti,
height, F is tbe liquid nutrient flow rate (cm31h), N, is flux of substrate into tbe bi,
(mg S/cm2 h), a is tbe biofilm or support partiele surface area per unit reactor
(cm2/cm3), A is tbe cross-sectional area of tbe bed (cm2), and dz is tbe differential
of an element of tbe column (cm), Substituting eq. 9.54 into eq. 9.69 yields the foJ
equation:

9.5. SOUD-STATE FERMENTATIONS
277
lneubation
Initial
moisture
Time
Temp.
Furtber
(%)
(h)
(oe)
processing
45
72
30
Yes
35
44
30
Yes
40
22
32
No
36
Yes
30
40
30
Tune
(min)
110
100
100
Temp.
(oe)
HI~rmal processing
Common
substrate
SOlid-State Fermentations
TABLE 9.3 SomeTraditional Food Fermentations
'T'L_. _
Prirnary
Product
genus
Soy sauce
Aspergillus
Miso
Aspergillus
Tempeh
Rhizopus
Hamanatto
Aspergillus
Sufu
SoyDean,
wheat
Riee,
soybean
Soybean
Soybean,
wheat
Actinomucor Tofu 100 10 74 15 Yes
With permission, from R. E. Midgett, in A. L. Demain and N. A. Solomon, eds., Manual oj Industrial Micro-
biology and Biotechnology, ACS Publieations. Washington, O.c.. 1986.
The major industrial use of the koji process is for the production of enzymes by
fungal species. Fungal amylases are produced by SSF of wheat bran by A. oryzae in a
rotating-drum fermenter. Wheat bran is pretreated with formaldehyde, and the initial pH
of the bran is adjusted to pH = 3.5 to 4.0 to reduce the chance of contamination. Usually,
perforated pans, rotating drums, or packed beds with air ventilation are used. A typical
rotary-drum type of koji fermenter is depicted in Fig. 9.16. Enzymes other than amylases,
such as cellulase, pectinase, protease, and lipases, can also be produced by koji fermenta-
tions. Trichoderma viride species have been used for the production of cellulases from
wheat bran in a rotary-tray fermenter.
Some secondary metabolites, such as antibacterial agents, are produced by Rhizopus
and Actinomucor species in some koji processes. Certain mycotoxins, such as aflatoxins,
were produced by SSF of rice (40% moisture) by A. parasiticus. Ochratoxins were also
produced by Aspergillus species on wheat in a rotary-drum koji fermenter. Microbial
degradation of Iignocellulosics can also be accomplished by soIid-state fermentations for
waste-treatrnent purposes or in biopulping of wood chips for use in paper manufacture.
Spores from some molds have found use as insecticides. Proper spore formation is diffi-
cult to obtain in submerged culture, and SSF must be used.
Major process variables in SSF systems are moisture content (water activity), inocu-
lum density, temperature, pH, particle size, and aeration/agitation. Optirnization of these
parameters to maximize product yield and rate of product formation is the key in SSF
systems and depends on the substrate and organism used. Most natural substrates
(e.g., grains) require pretreatrnent to malce the physical structure of substrates more sus-
ceptible to myceIial penetration and utilization. Solid substrates are usually treated with
antimicrobial agents, such as formaldehyde, and are steamed in an autoclave. Nutrient
media addition, pH adjustrnent, and the adjustment of moisture level are realized before
inoculation of the fermentation mash. Koji fermentations are usually realized in a
controlled-hurnidity air environment with air ventilation and agitation. Many soIid-state
mycelial fermentations are shear sensitive due to disruption of the myceIia at high
Sec. 9.5
Bioreaetors for Suspension and Immobilized Cultures
276
Solid-state fermentations (SSF) are fermentations of solid substrates at low moisture lev-
els or water activities. The water content of a typical submerged fermentation is more than
95%. The water content of a solid mash in SSF often varies between 40% and 80%.
Solid-state fermentations are usually used for the fermentation of agricultural products or
foods, such as rice, wheat, barley, com, and soybeans. The unique characteristic of
solid-state fermentations is operation at low moisture levels, which provides a selective
environment for the growth of mycelial organisms, such as molds. In fact, most solid-state
fermentations are mold fermentations producing extracellular enzymes on moist agricul-
tural substrates. Since bacteria and-yeasts cannot tolerate low moisture levels (water activ-
ities), the chances of contarnination of fermentation media by bacteria or yeast are greatly
reduced in SSF. A1though most SSFs are mold fermentations, SSFs based on bacteria and
yeast operating at relatively high moisture levels (75% to 90%) are also used. Solid-state
fermentations are used widely in Asia for food products, such as tempeh, rniso, or soy
sauce fermentations, and also for enzyme production.
The major advantages of SSFs over submerged fermentation systems are (I) the
small volume of fermentation mash or reactor volume, resulting in lower capital and oper·
ating costs, (2) a lower chance of contarnination due to low moisture levels, (3) easy prod"
uct separation, (4) energy efficiency, and (5) the allowing of the development of fully
differentiated structures, which is critical in some cases to product formation. The major
disadvantage of SSFs is the heterogeneous nature of the media due to poor mixing cbar:
teristics, which results in control problems (pH, DO, temperature) within the fermentati()]
mash. To elirninate these control problems, fermentation media are usually rnixed eithl
continuously or interrnittently. For large fermentation mash volumes, the concentratiOl
gradients may not be elirninated at low agitation speeds, and mycelial cells may be d:
aged at high agitation speeds. Usually, a rotating-drum fermenter is used for SSF syste.
and the rotational speed needs to be optimized for the best performance.
Solid-substrate fermentations imply a more general method of fermentations .
which moisture content may not need to be low, but the substrate is in the form of s
merged solid particles in liquid media. Bacterial ore leaching (Le., growth and microb'
oxidation on surfaces of rnineral sulfide particles) or fermentation of rice in a pac
colurnn with circulating liquid media are examples of solid-substrate fermentati
Solid-state (or solid-phase) fermentations are a special form of solid-substrate ferme'
tions for which the substrate is solid and the moisture level is low.
The koji process is an SSF system that employs molds (Aspergillus, Rhizop,
growing on grains or foods (wheat, rice, soybean). A typical SSF process involves
stages. The first and the primary stage is an aerobic, fungal, solid-state fermentation
grains cal1ed the koji. The second stage is an anaerobic submerged fermentation wi'
mixed bacterial culture called the moromi. The products listed in Table 9.3 are the
ucts of aerobic SSF, the koji process. Fermentation in the second stage (mororni) may
realized by using the natural flora, or, usual1y, with externally added bacteria and y'
Some strains of Saccharomyces, Torulopsis, and Pediococcus are used as flavor produ-
in soy sauce manufacture. The moromi is usually fermented for 8 to 12 months. How,
the processing time can be reduced to 6 months by temperature profi1ing. The final
uct is pressed to recover the liquid soy sauce and is pasteurized, filtered, and bottled.

279
o
WHEAT BRAN
Summary
volume, low-product-value processes (e.g., waste treatrnent and fuel-grade ethanol pro-
duction). Multistage continuous systems improve the potential usefulness of continuous
processes for the production of secondary metabolites and for the use of genetically unsta-
ble cells. The fed-batch system is widely used in commercial plants and combines the fea-
tures of continuous culture and batch that allow the manufacturer to maintain flexibility
and ease of intervention. The perfusion system is another option that is particularly attrac-
tive for animal cells.
lmmobilized cell systems offer a number of potential processing advantages, and
the commercialization of such systems is proceeding rapidly where cell culture is expen-
sive and difficult (e.g., animal cell tissue culture). Physical entrapment or encapsulation is
used in most cases, although adsorption onto surfaces or covalent binding of cells to sur-
faces is possible.
In some cases, self-immobilization on surfaces is possible and a biofilm is formed.
Biofilm reactors can apply to tissue culture, mold, and bacterial systems. Biofilm-based
reactors are very important in waste-treatrnent applications and in natural ecosystems. The
analysis of immobilized cell reactors is analogous to that for immobiIized enzyme reac-
tors except for the feature of biocatalyst replication.
Solid-state fermentations share sorne characteristics with immobilized cell systems,
but differ in that no discernible liquid is present. SSFs have found important uses in the
production of sorne traditional fermented foods and may have use in upgrading agricul-
tural or forest materials and in the production of mold products requiring full mold differ-
entiation.
Figure 9.17. Rotary, automatic koji-making apparatus. Tbe apparatus has a two-storied
chamber. Each chamber has a large rotary tray on which wheat bran is heaped evenly.
After inoculated fungus has grown sufficiently, solid culture is transferred by a screw con-
veyor to the lower rotary-tray hopper. (With permission, from N. Toyama. Biotechnol.
Bioeng. Symp., Vol. 6, pp. 207-219,1976, John Wiley & Sons, Inc .• NewYork.)
ROTARY CHAMBER FOR SOLID CULTURE
IMNHOLE
INLET DUCT
::~::::~:::::::=;
HEATEIl

'i~'l',.
.•••
1·1'i Yi ",S
278
AU1O'AATIC SOLlO CULTUIlE APf'ARATUS (IIOTARY PIIOCOSI
Figure 9.16. Rotary-drum type of koji-making apparatus used for rice solids culture by
A. oryzae. AU operations (washing, cooking, inoculation, loosening of solids, water spray-
ing, cooling, air circulation, fi1ling, and exhausting) can be done in this apparatus. (With
permission, from N. Toyama, Biotechnol. Bioeng. Symp., Vol. 6, pp. 207-219, 1976,John
Wiley & Sons, Inc., New York.)
agitationlrotation speeds. At low agitation rates, oxygen transfer and CO2 evolution ral
become limiting, Therefore, an optimal range of agitation rate or rotation speed needs
be determined. Similarly, there is a minimum level of moisture content (-30% by wei; ..
below which microbial activity is inhibited. At high moisture levels (>60%), solid SI
strates become sticky and form large aggregates. Moreover, moisture level affects
metabolic activities of cells. Optimal moisture level needs to be determined experi
tally for each cell-substrate system. For most of the koji processes, the optimal moisl
level is about 40% ± 5%. partic1e size should be small enough to avoid any oxygen
exchange or other nutrient transport limitations. Porosity of the partic1es can be imprl
by pretreatrnent to provide a larger intrapartic1e surface to volume ratio for mk
action.
Most of the SSF processes are realized using a rotary-tray type of reactor in a
perature- and humidity-controlled chamber where controlled-humidity air is circ
through stacked beds of trays containing fermented solids. Figure 9.17 depi
rotary-tray chamber for koji fermentations. Rotary-drum fermenters are used le:
quently because of the shear sensitivity of mycelial cells.
OI1TLET DUCT
9.6. SUMMARY
Bioreactors using suspended cells can be operated in many modes intermediate be
batch reactor and a single-stage chemostat. Although a chemostat has potential p
ity advantages for primary products, considerations of genetic instability, process
ity, and proces s reliability have greatly limited the use of chemostat units. The
recyc1e with a CSTR increases volumetric productivity and has found use.

PROBLEMS
SUGGESTIONS FOR FURTHER READING
9.1. Consider a 1000-1 CSTR in which biomass is being produced with glucose a$
The ITUcrobialsystem follows a Monod relationship with J.lm = 0.4 h-1, Ks = 1.5
281
Tlme X
P
dXJdt
dPldt
h g/lg/l
g/l-h
g/l-h
O
0.3
<0.01
3 1.0
<0.01
0.30
6 2.3
<0.01
0.55
8 4.00.010
1.0
0.005
9 5.1
0.025
1.3
0.010
10 6.5
0.060
1.4
0.045
10.5 7.0
-
1.4
11 7.4
0.10
0.60
0.059
12 7.7
0.17
0.20
0.072
13 7.8
0.26
0.02
0.105
14 -
0.36
-
0.130
15 8.0
0.47
-O
0.087
16 8.0
0.54
-O
0.042
17 -
0.58
-
0.021
18 -
0.60
-
0.005
J.lm = 0.3 h-l, g dw cells
Ks = 0.1 g/l, YXlS = 0.4 _
g glucose
a. Determine cell and glucose concentrations in the effluent of the fust stage.
b. Assume that growth is negligible in the second stage and the specific rate of product for-
mation is qp = 0.02 g PIg cell h, and YPIS = 0.6 g Plg S. Determine the product and sub-
strate concentrations in the effluent of the second reactor.
Consider the following batch growth data:
ally high value), and the yield factor YX1S = 0.5 g biomass/g substrate consumed. If normal
operation is with a sterile feed containing 10 gII g1ucose at a rate of 100 1Ih:
a. What is the specific biomass production rate CgII-h)at steady state?
b. If recycle is usOOwith a recycle strearn of 10 IIh and a recycle biomass concentration five
times as large as thal in the reactor exit, what would be the new specific biomass produc-
tion rate?
c. Explain any difference between the values found in parts a and b.
9.2. In a two-stage chemostat system, the volumes of the fust and second reactors are V] = 500 I
and V2 = 300 1, respectively. The first reactor is usOOfor biomass production and the second is
for a secondary metabolite formation. The feed flow rate to the first reactor is F = 100 1Ih,
and the glucose concentration in the feed is S = 5.0 g/l. Use the fOllowing constants for the
cells.
9.3.
You have available three tanks of different volumes: 900, 600, and 300 I. Given a flow rate of
100 1Ih, what configuration oftanks would maxiITUzeproduct formation?
9.4. Penicillin is produced by P. chrysogenum in a fed-batch culture with the intermittent addition
of glucose solution to the culture medium. The initial culture volume at quasi-steady state is
Vo = 500 I, and glucose-containing nutrient solution is added with a flow rate of F = 50 1Ih.
G1ucose concentration in the feed solution and initial cell concentration are So = 300 gII and
Xo = 20 gII, respectively. The kinetic and yield coefficients of the organism are J.lm = 0.2 h-l,
Ks = 0.5 gII, and YX1S = 0.3 g dw/g glucose.
Bioreactors for Suspension and Immobilized Cultu
280
The reactor options described in this chapter are many. The best choice of reactor
systems wiU ultimately be determined by the choice of biocatalyst and the requirements
for product recovery and purification.
AlBA, S., A. E. HuMPHREY,ANDN. F. MILLlS,Biochemical Engineering, 2d ed., AcadeITUc
NewYork,1973.
ATKINSON,
B., ANDF. MAVITUNA,
Blóchemical Engineering and Biotechnology Handbook, 2d
Stockton Press, NewYork, 1991.
BAlLEY,J. E., ANDD. F. OLLlS,Biochemical Engineering Fundamentals, 2d 00., McGraw-Hill
CO., NewYork, 1986.
BLANCH,
H. W., ANDD. S. CLARK,Biochemical Engineering, Marcel Dekker, Inc., New York, 199,
CHARACKLlS,
W. G., R. BAKKE,ANDM. G. TRULEAR,1991, "Fundarnental Considerations ofE
Film Systems," in M. Moo-Young, 00., Comprehensive Biotechnology, Vol. 4, pp. 945-'
1985.
CHlBATA,
I., T. TosA, ANDT. SATO,"Methods of Cell Immobilization," in Manual of 1ndustria/
crobiology and Biotechnology, A. L. Demain and N. A. Solomon, eds., American Society
Microbiology, Washington, DC, pp. 217-229, 1986.
DEGoOUER,C. D., W. A. M. BAKKER,H. H. BEEFTINK,
ANDJ. 1'RAMPER,
Bioreactors in Se' .
Overview of Design ProcOOures and Practical Appllcations, Enz;yme Microbiol Technol,
202-219,1996.
KARGI,F., ANDM. MOO-YoUNG,"Transport Phenomena in Bioprocesses," in M. Moo-Ya
Comprehensive Biotechnology, Vol. 2, Pergarnon Press, Elmsford, NY, pp. 5-55,1985.
KLElN,J., ANDK. D. VORLOP,"Irnmobi1ization Techniques: Cells," in M. Moo- Young, 00., C.
hensive Biotechnology, Vol. 2, Pergarnon Press, Elmsford, NY, pp. 203-334, 1985.
MERCILLE,S., M. JOHNSON,
S. LAUTH!ER,
A. A. KAMEN,ANDB. MASSIA,Understanding F:
LiITUtthe Productivity of Suspension-BasOO Perfusion Cultures Operated at High M
newal Rates, Biotechnology Bioengineering, 67: 435-450, 2000.
MIDGETI,R. E., "Solld State Ferrnentations," in A. L. Demain and N. A. Solomon, Manual
trial Microbiology and Technology, American Society for Microbiology, Washin:
pp. 66-83,1986.
Moo-YoUNG, M., Bioreactor 1mmobilized Enzymes and Cells: Fundamentals and APJ
Elsevier Science Publlshing, Inc., New York, 1988.
SCHROEDER,
E. D., Water and Wastewater Treatment, McGraw-Hill Book Co., New York,
WANG,D. I. C., ANDaTHERS,Fermentation and Enzyme Technology. John Wiley & Sons,
WEBSTER,I. A., M. L. SHULER,ANDP. RONY.The Whole Cell Hollow Fiber Reactor:
Factors, Biotech. Bioeng. 21: 1725-1748, 1979.

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