This document summarizes a study that evaluated gas-liquid mass transfer and scale up of a biotransformation process from shake flasks to a 5 L stirred tank bioreactor. The study used a yeast isolate capable of converting benzaldehyde to l-phenyl acetyl carbinol (l-PAC) in both growth and biotransformation media. Experiments characterized gas holdup, power input, and mass transfer coefficient (KLa) for the media using different impeller combinations and operating parameters. Results were used to optimize conditions for maximal cell growth and l-PAC production in the bioreactor, which were then compared to shake flask studies. Correlations developed could predict mass transfer and be applied to scale up the process.

1. Biochemical Engineering Journal 8 (2001) 19–29
Scale-up of biotransformation process in stirred tank reactor using
dual impeller bioreactor
V.B. Shuklaa
, U. Parasu Veerab
, P.R. Kulkarnia
, A.B. Panditb,∗
a Department of Chemical Technology, Food and Fermentation Technology Division, University of Mumbai (UDCT), Matunga, Mumbai 400 019, India
b Department of Chemical Technology, Division of Chemical Engineering, University of Mumbai (UDCT), Matunga, Mumbai 400 019, India
Received 29 April 1999; accepted 14 November 2000
Abstract
The gas–liquid mass transfer coefficient KLa in the fermenter is a strong function of mode of energy dissipation and physico-chemical
properties of the liquid media. A combination of disc turbine (DT) and pitched blade turbine down flow (PTD) impellers has been tested
in laboratory bioreactor for gas hold-up and gas–liquid mass transfer performance for the growth and biotransformation medium for an
yeast isolate VS1 capable of biotransforming benzaldehyde to l-phenyl acetyl carbinol (l-PAC) and compared with those in water.
Correlations have been developed for the prediction of the fractional gas hold-up and gas–liquid mass transfer coefficient for the
above media. The mass transfer coefficient and respiration rate have been determined in the shake flask for the growth as well as for
biotransformation medium. These results, then have been used to optimize the operating parameters (impeller speed and aeration) for
growth and biotransformation in a laboratory bioreactor. The comparison of cell mass production and l-PAC production in the bioreactor
has been done with that obtained in shake flask studies. © 2001 Elsevier Science B.V. All rights reserved.
Keywords: Gas–liquid mass transfer; Dual impeller; Bioreactor; Fermentation; Oxygen solubilization; Scale-up; Respiration rate; l-phenyl acetyl carbinol
1. Introduction
Growth of organism like yeast, in a medium rich in nu-
trient requires large aeration for complete utilizations for
building of cell mass. The catalytic activities of the organ-
ism are fully utilized if the oxygen level can be maintained
at a particular concentration in the immediate vicinity of the
cells [1]. At the same time shear due to aeration and agita-
tion should be kept at a reasonable level as the higher shear
rates may physically damage the cells, thereby affecting the
microorganisms.
In the case of growth of yeast cells nutrient rich medium
can be fully utilized to give cell mass if dissolved oxygen
concentration can be maintained at a desired level in the
direct vicinity of the cells. During biotransformation of ben-
zaldehyde to l-phenyl acetyl carbinol (l-PAC) by yeast cells
the level of oxygen should be such that it should maintain
organisms in the active form to maximize the production
of l-PAC and reduce the production of other by-products.
Continuous stirred tank reactor (CSTR), is commonly used
in many of the bioprocess as it allows efficient contacting
of three phases, i.e. gas, liquid medium and solid cells.
∗ Corresponding author. Tel.: +91-22-414-5616; fax: +91-22-414-5614.
E-mail address: abp@udct.ernet.in (A.B. Pandit).
In the case of CSTR, the aspect ratio is maintained more
than 1, to ensure high residence time of gas phase, increase
the transfer efficiency and to ensure less power input on
introduction of gas, uniform power dissipation. The shear
can be reduced by operating the impellers at lower speed.
For the same volume of liquid, multiple impellers on a
single shaft with appropriate combination and spacing are
being suggested as optimum [2]. In such cases mixing and
mass transfer are dependent on the flow rate of gas, type
of agitator and its speed and properties of liquids. Power
consumption per impeller decreases with an increase in the
number of impellers [2] and this increases the uniformity
of energy dissipation. A combination of DT–PTD has been
reported to give highest values of fractional gas hold-up
and of gas–liquid mass transfer coefficient with least power
input for air–water system by Arjunwadkar et al. [3].
In the case of biotransformation of benzaldehyde to l-PAC
(which is an intermediate of various ␣- and ␤-adrenergic
drugs like ephedrine and pseudoephedrine) growing of
the cell mass and biotransformation of benzaldehyde un-
der regulatory conditions to obtain maximal cell mass
and l-PAC are the main aims. Under normal fermentation
conditions quantitative conversion of benzaldehyde into
l-PAC has never been achieved [5,6]. Formation of l-PAC
has been reported to be associated with the formation of
1369-703X/01/$ – see front matter © 2001 Elsevier Science B.V. All rights reserved.
PII: S1369-703X(00)00130-3

2. 20 V.B. Shukla et al. / Biochemical Engineering Journal 8 (2001) 19–29
Nomenclature
a interfacial area (m2 m−3)
Cfinal final oxygen concentration (kg m−3)
Cin initial oxygen concentration (kg m−3)
C∗ saturation oxygen concentration
(kg m−3)
D diameter of the impeller (m)
DT disc turbine
H height of dispersion (m)
HL height of clear liquid height (m)
KLa mass transfer coefficient (s−1)
N impeller speed (r s−1 or rps)
NQ gas flow number
P0 power consumption under non-aerated
conditions (W)
PG power consumption under gassed
conditions (W)
PTD pitched blade turbine down flow
QG gas flow rate (m3 s−1)
R respiration rate (mg s−1)
t time duration of aeration to reach initial
to final oxygen concentration
V volume of liquid in bioreactor (l)
VVM volume of gas per volume of aerated
liquid (min−1) per minute
Greek symbols
εG fractional gas hold-up
by-products like benzyl alcohol, 1-phenyl-1,2-propanediol
(PAC-diol) and benzoic acid. The present study aims at the
scale-up of the biotransformation of benzaldehyde to l-PAC
from shake flask to 5 l bioreactor with KLa as a scale-up
criteria.
2. Experimental, materials and techniques
For shake flask and bioreactor studies, growth medium
having composition of glucose 3%, peptone 2%, and yeast
extract 1% (pH 5.5, optimized value) and biotransformation
medium having composition of glucose 3%, peptone 0.6%
(pH 4.5, optimized value) were used. In the case of shake
flask studies growth of yeast isolate VS1 (identified as
Saccharomyces cerevisiae) having capacity to transform
benzaldehyde to l-PAC was carried out. For growth, 1 ml
of inoculum containing 1 × 106 cells were inoculated into
9 ml of growth medium and incubated on a rotary shaker
at 30 ± 2◦C at 240 rpm for 24 h. The culture so obtained
was used as an inoculum for 100 ml of the growth medium,
allowed to grow for 24 h under similar conditions. It was
centrifuged with Beckman J2-MC centrifuge operated at
relative centrifugation field of 17000g and at 10,000 rpm
for 15 min after chilling it to 15◦C. The biomass so ob-
tained was used as an inoculum for the 100 ml of the
biotransformation medium. The cells were allowed to adapt
to biotransformation medium for 1 h on the rotary shaker
at 240 rpm then added 0.6% (w/v) of benzaldehyde and
to 0.6% (v/v) of acetaldehyde (30–35%) in the medium.
The biotransformation was then allowed to take place on
a rotary shaker under same conditions and the samples
were drawn at regular time intervals (20 min each). The
transformed medium was centrifuged at 10,000 rpm for
15 min. The supernatant so obtained was extracted thrice
with equal volume of diethyl ether. The extract was dried
over anhydrous Na2SO4 and concentrated under vacuum
at 30◦C temperature. The concentrated extract thus ob-
tained was subjected to gas-chromatographic analysis.
The conditions used for GC were as follows: GC model
Chemito-8510 with Oracle-1 computing integrator, col-
umn used was 5%OV-17 4 m long, injector temperature
of 250◦C and detector temperature of (FID) 250◦C, col-
umn programming was as follows: 75◦C for 3 min then
10◦C min−1 up to 250◦C and holding for 5 min. Retention
times of benzaldehyde, benzyl alcohol, l-PAC, PAC-diol
were 11, 13, 17 and 18.5 min, respectively. The concentra-
tions of these compounds were determined using peak area
method.
The laboratory scale fermenter (Chemap AG) as described
in Fig. 1 has been used in this study. Details about the di-
mensions of experimental set-up are given in Fig. 2. Liq-
uid height used was 0.21 m (HC/D = 1.167) and liquid
volume used was 5.125 l. Air was sparged through a pipe
sparger placed below the bottom impeller. The gas flow rates
used were varied from 1.5 to 8 l min−1 for growth (VVM
range 0.293–1.56) and 1.5–3 l min−1 for biotransformation
medium (VVM range 0.293–0.586). Impeller speed was var-
ied from 50 to 300 rpm for mass transfer studies. Impellers
used in the study were disc turbine (DT) as a lower impeller
and pitched blade down flow turbine (PTD) as an upper im-
peller.
An AC/DC current transducer (Clamp tester) (Appa 32;
Appa Technology, Taiwan) working on Hall’s effect was
used for power measurement by connecting it to a digital
multimeter (MECO model 9A; MECO Instrument, Seweree
Mumbai) for actual read out. Fractional gas hold-up was
measured by visual observation. Power consumption was
found to be reproducible within ±5%. The oxygen concen-
tration in liquid phase was measured using polarographic
probe (Russel, UK) with fast responding PTFE membrane.
The microbial media ingredients were purchased from M/S
Hi Media Laboratories, Mumbai, India.
In the case of shake flask as well as bioreactor studies, the
mass transfer coefficient (KLa) was calculated by measur-
ing the rate of oxygen transfer in nitrogen (IOLAR grade)
purged medium. A typical response curve for oxygen con-
centration against time for shake flask and bioreactor (at
different rpm and aeration rates) for growth and biotrans-
formation medium were drawn. The KLa was calculated as

3. V.B. Shukla et al. / Biochemical Engineering Journal 8 (2001) 19–29 21
Fig. 1. Schematic representation of fermenter vessel and associated set-up. 1, Pressure holding valve; 2, exit air filter; 3, harvest port; 4, condensation
coil; 5, glass U tube for hold-up measurement; 6, temperature sensor (to control panel); 7, sample port; 8, stem supply for sample port; 9, tube sparger;
10, shaft; 11, impeller; 12, notch for impeller fitting; 13, hollow baffle; 14, antifoam probe (to control panel); 15, air inlet filter; 16, water for cooling;
17, protection grid; 18, bottom plate; 19, top plate.
follows [4]:
KLa =
1
t
ln
C∗ − Cfinal
C∗ − Cin
(1)
In the case of bioreactor for growth of organism 24 h old
culture (2 ml × 110 ml grown as mentioned above) was
Fig. 2. Geometrical dimensions of bioreactor set-up.
inoculated into 5 l of sterile growth medium. The growth
was monitored in terms of absorbance at 660 nm (using Hi-
tachi spectrophotometer model 2001; Hitachi Instruments,
Japan). After 24 h growth of organism the medium was sub-
jected to centrifugation at 10,000 rpm and 15◦C for 15 min
(using Backman centrifuge model J2-MC). The cell mass so

4. 22 V.B. Shukla et al. / Biochemical Engineering Journal 8 (2001) 19–29
obtained was inoculated into 5 l of biotransformation
medium by using peristaltic pump into the reactor under
aseptic conditions. After 1 h adaptation of cell mass in the
biotransformation medium, 0.6% (w/v) of benzaldehyde
and 0.6% (v/v) of acetaldehyde (30–35%) were added asep-
tically. The samples were withdrawn at regular intervals of
20 min and subjected to extraction, concentration and GC
analysis as mentioned previously.
3. Results and discussion
3.1. Power dissipation
The relative power input PG/P0, were measured for the
given liquid phases against gas flow number (constant flow
rate and variable impeller speed). Depending on the flow
regime, i.e. flooding, loading, dispersion and re-circulating,
trends were similar for both the growth medium and bio-
transformation medium (Fig. 3) for the DT–PTD impeller
combination. In the case of other impellers studied by Ar-
junwadkar et al [3], the lower impeller passes through all
the regimes whereas, the upper impeller remains in disper-
sion or re-circulating regime. It can be seen from Fig. 3, that
flow number is high for lower impeller speeds, which causes
Fig. 3. Relative power input against gas flow number for different media at QG = 1.5 l min−1
and variable N. (᭜) Water, (᭿) growth medium and (᭡)
biotransformation medium.
flooding and loading, where the relative power input is
also lower. As the impeller speed increases, the operational
regime changes to dispersion or re-circulating. Therefore,
the bubbles are dragged in the down flow region of the agi-
tator. Because of this, the gas hold-up in the agitator region
increases which results in a continuous decrease in power
consumption. The given impeller combination (DT–PTD)
is observed to handle larger quantities of gas for the lowest
PG/P0. As the gas flow rate increases (VVM, volume of
gas per volume of unaerated liquid per minute) the value
of PG/P0 decreases continuously, because of the higher
hold-up. Fig. 4 shows the variation of power input with
the VVM for water, growth medium and biotransformation
medium. According to the controlling energy input criteria,
the effect of impeller action is much more at lower gas flow
rates than at higher gas flow rates. As a consequence of this
criteria, the power input drops to a higher extent (Fig. 4)
initially (for air–water) and with an increase in the gas flow
rates it then marginally decreases up to VVM 1.6, studied
in this work. For growth and biotransformation medium,
the drop in power input is marginal as the physico-chemical
properties of these liquids affect the bubble coalescence
behavior leading to this kind of variation.
This observation is consistent with the observation
made in the variation of PG/P0 for viscous liquids [7]. On

5. V.B. Shukla et al. / Biochemical Engineering Journal 8 (2001) 19–29 23
Fig. 4. Relative power input against gas flow number for different media at constant rpm (N = 250) and variable QG. (᭜) Water, (᭿) growth medium
and (᭡) biotransformation medium.
introduction of the gas in a viscous liquid, initially there
a tendency for the gas to get accumulated behind the im-
peller blade and form a cavity. The size of this cavity,
which decides the reduction in PG/P0 is known to be a
function of impeller speed alone and not of the volumetric
gas flow rate. Thus, the initial sharp decrease in PG/P0 is
possibly due to the formation of these cavities, which does
not change with an increase in QG further. The continuous
decrease in PG/P0 though marginal could possibly due to
the accumulation of the small bubbles known to occur in
viscous liquid altering the dispersion density. The behavior
for the air–water system is consistent with the observation
reported in the literature [3].
3.2. Fractional gas hold-up
Fractional gas hold-up is an important design parameter
for gas–liquid and gas–liquid–solid systems. Gas hold-up
determines the required size of the reactor and the interfa-
cial area which in turns decides the gas–liquid mass transfer
coefficient [8]. The impeller combination used (DT–PTD)
in this study is based on the work done by Arjunwadkar
et al. [3], which has given highest gas hold-up in air–water
system. Gas hold-up has been studied for three liquids, viz.
air–water, growth medium and biotransformation medium.
The value of the gas hold-up was estimated using the fol-
lowing equation:
εG =
H − HL
H
(2)
The fractional gas hold-up was found to be the highest in
the air–water system (Fig. 5a) than in other two systems as
shown in Fig. 5b and c. The gas hold-up in growth medium is
higher than the biotransformation medium. This may be due
to the variation in the medium composition and variation in
their physico-chemical properties. The hold-up for biotrans-
formation medium was measured only for the flow rate of
1.5 and 3 l min−1. For the higher air flow rates, stable foam
formation was observed making hold-up measurements dif-
ficult. The reduction in the gas fraction in both these media
over the air–water system may be due to the presence of the
glucose (3%) along with other compounds, which makes the
solution slightly viscous, promoting the bubble coalescing
tendency increasing the average bubble size, reducing the
hold-up values. This observation is also consistent with the
previous observation of the variation in PG/P0 as the extent
of reduction in PG/P0 for biomedia was always less than
that for air–water system indicating the influence of hold-up
on the same.
3.3. Oxygen mass transfer coefficient
The mass transfer coefficient (KLa) was measured
for the given impeller combination as per the procedure

6. 24 V.B. Shukla et al. / Biochemical Engineering Journal 8 (2001) 19–29
Fig. 5. (a) Effect of power on hold-up at different aeration for air–water system. (᭜) 3 l min−1, (᭿) 5 l min−1, (᭡) 8 l min−1, (᭹) 10 l min−1. (b) Effect
of power on hold-up at different aeration for growth medium. (᭜) 1.5 l min−1, (᭿) 3 l min−1, (᭡) 5 l min−1, (×) 8 l min−1. (c) Effect of power on
hold-up at different aeration for biotransformation medium. (᭜) 1.5 l min−1, (᭿) 3 l min−1.
described earlier. It was found that the KLa was highest in
the air–water system than in growth and biotransformation
medium as shown in Fig. 6a–c. This has been discussed
by Arjunwadker et al. [4]. The highly protenaceous nature
of the growth medium offers extra resistance for the mass
transfer at the interface due to the interfacial absorption of
the protein molecules. The viscosity also increases due to
the presence of the glucose in the medium thereby decreas-
ing the mass transfer coefficient. KLa in biotransformation
medium is less than the growth medium as it does not contain
the nutrients in the form of salts, as the aqueous solution of
salt acts as a coalescence inhibiting medium. The variation
in the KLa is consistent with the variation in the εG reported
earlier.
3.4. Correlations
Correlations for gas hold-up and mass transfer coeffi-
cient KLa with operating variables (power consumption and
superficial gas velocity) have been developed for growth
and biotransformation medium for this dual impeller sys-
tem.

7. V.B. Shukla et al. / Biochemical Engineering Journal 8 (2001) 19–29 25
Fig. 6. (a) Effect of power input on mass transfer coefficient in air–water system (VG = 0.00196). (b) Effect of power input on mass transfer coefficient
in growth medium (VG = 0.00196). (c) Effect of power input on mass transfer coefficient in biotransformation medium (VG = 0.00196).
3.4.1. Gas hold-up
The correlation of the type, proposed by Abradi et al.
[9] for dual impeller system has been proposed for growth
medium and is given as follows:
εG = 0.794
PG
V
(1 − εG)
0.482
VG
0.601
,
R2
= 89% (3)
Similarly for Biotransformation medium,
εG = 0.09
PG
V
(1 − εG)
1.171
VG
0.5624
,
R2
= 90% (4)
The correlation for air–water system proposed by Arjun-
wadkar et al. [3] for the same impeller combination was
found to be valid and is reproduced here as,
εG = 0.064
PG
V
(1 − εG)
0.40
VG
0.50
, R2
= 80.5%

8. 26 V.B. Shukla et al. / Biochemical Engineering Journal 8 (2001) 19–29
The variation in the constants was found in all the above
three equations. This variation may be due to the dif-
ference in the composition of the growth and biotrans-
formation medium. The exponent of the first term in Eq. (4)
is quite different from that of Eq. (3). Correlation for
biotransformation was fitted for the air flow rates of 1.5 and
3 l min−1. For the higher flow rates stable foam formation
was observed and this must have increased the fractional
gas hold-up substantially, though due to the visual measure-
ment technique, an error might have introduced resulting
under estimation of hold-up. This could be a reason for a
very strong dependence of hold-up on power dissipated per
volume (exponent 1.171). Except Eq. (4), all the constants
and exponents obtained are reasonably consistent with the
reported literature [9].
3.4.2. Mass transfer coefficient
Arjunwadkar et al. [4] established the consistency with
the published work for the DT–PTD impeller combina-
tion for the air–water system with the correlation of the
type
KLa = a
PG
V
b
VG
c
(5)
The constants in the above correlation for air–water system
are a = 2.04 × 10−3, b = 0.68 and c = 0.58 with cor-
relation coefficient of 96.4%. On the similar lines the cor-
relations have been developed for the growth medium and
biotransformation medium.
For growth medium,
KLa = 0.024
PG
V
0.725
V 0.892
G , R2
= 82% (6)
Similarly for biotransformation medium,
KLa = 6.99 × 10−6 PG
V
1.14
V 0.365
G , R2
= 87% (7)
The variation in the constants may be due to the difference
in the chemical composition of the mediums and viscosity of
the broths. Again, as the KLa measurements involved no vi-
sual estimation of the gas hold-up and still the exponent over
(PG/V) is high, i.e. 1.14 as against 0.68 (for air–water) 0.725
(for growth medium). In Section 3.4.1, similarly a high ex-
ponents over {PG/V (1 − εG)} term has been reported while
correlating fractional gas hold-up. This fact, along with the
observation of considerable foaming, possibly explain the
strong dependence of εG and KLa on (PG/V). The change in
the nature of the media from coalescing to non-coalescing
type is possibly responsible for such exponents.
3.5. Scale-up
The power dissipation level in the shake flask is indeed
different from that of CSTR. It cannot be estimated inde-
pendently like in the case of the impeller and hence, the
observed/measured KLa values obtained in the shaker flask
are used to get equivalent values in CSTR and no attempt
has been made to develop a correlation for KLa for shake
flask like the one developed in the case of CSTR.
The production of l-PAC was scaled up to the 5 l biore-
actor from shake flask with same KLa as the criteria. First
the KLa was estimated in the shake flask at different times
during both, growth of the yeast and the biotransformation
process. Initially the KLa in the shake flask were estimated
by keeping them on shaker at 240 rpm and monitoring the
dissolved oxygen concentration at regular intervals for both
the media at different stages of growth. Before the estimation
of the dissolved oxygen the broths were purged with nitro-
gen to bring down the DO to the minimum level. This was
treated as zero time, DO was then measured with time, at
regular intervals to estimate KLa during the various phases of
growth and biotransformation. Similar values of mass trans-
fer coefficients (KLa) were maintained in the 5 l bioreactor
for the growth and biotransformation using the correlations
developed earlier (Eqs. (6) and (7)) and by the manipulation
of power input and aeration rates. For the estimation of the
power input and aeration rates, the oxygen respiration rates
were estimated for growth and biotransformation, so that
the desired DO level could be maintained with gas-flow and
impeller speed combination giving a specific value of KLa.
3.6. Estimation of respiration rates
The rate of change of dissolved oxygen with respect to
time is very high immediately after inoculation, and it is
very nominal as the time progresses. For this reason the res-
piration rates was estimated for two different stages. This
total respiration rate was used for the estimation of required
net KLa (two stages, after inoculation and during equilib-
rium stage). The medium was saturated with the air till DO
remained unchanged with time, and after that the aeration
was stopped. After reaching the saturation point, the shake
flask was inoculated. The decrease in DO was monitored
continuously with time. There was a sharp decrease in DO
level initially for some time and then more or less remain
unchanged (Fig. 7a and b). Respiration rates were estimated
for both the parts of these curves. The slopes for these two
portions of the curve will give the respiration rate. The res-
piration rate can be estimated as
dc
dt
=
Cin − Cfinal
t
(8)
Slope of the graph for the two parts of the curve indicated by
R1 = (dc/dt)I and R2 = (dc/dt)II which are the slopes of
the first portion (AB) and second portion (BC) of the curve
(as shown in Fig. 7a and b for both media), where R1 is the
respiration rate immediately after inoculation of organism in
the media (curve AB) and R2 is the respiration rate (curve
BC). Thus the net respiration rate is given by
R = R1 + R2 (9)

9. V.B. Shukla et al. / Biochemical Engineering Journal 8 (2001) 19–29 27
Fig. 7. (a) Variation of dissolved oxygen with respect to time in growth medium. (b) Variation of dissolved oxygen with respect to time in biotransformation
medium.
Thus the net required KLa can be found out by equat ing
this respiration rate with the oxygen transfer rate as
R = KLa(C − C∗
) (10)
Thus, the required
KLa =
R1 + R2
C − C∗
(11)
which were used to decide PG/V and VG in the 5 l bioreactor
for different liquid media based on the correlations (Eq. (6)
and (7)).
3.7. Comparison of shake flask results with bioreactor
results
The growth in the shake flask and bioreactor have
been compared on the basis of wet cell mass weight

10. 28 V.B. Shukla et al. / Biochemical Engineering Journal 8 (2001) 19–29
Table 1
Comparison of product profile in shake flask and bioreactor at regular time intervals
Time (min) Shake flask (mg/100 ml) Bioreactor (mg/100 ml)
l-PAC Benzyl alcohol PAC-diol Residual benzaldehyde l-PAC Benzyl alcohol PAC-diol Residual benzaldehyde
20 194 191 0 172 244 80 0 206
40 220 214 0 153 361 174 0 84
60 270 242 0 121 411 259 0 15.8
80 304 289 0 51 412 286 2.0 11.8
100 310 305 0.2 25 426 288 4.5 1.2
120 330 345 0.9 3.6 458 304 6.6 0.8
140 330 346 3.2 – 448 301 9.6 0.7
160 – – – – 440 300 11.8 0.6
180 – – – – 437 300 14.8 0.5
540 – – – – 416 300 38.8 0.5
produced after 24 h. It was found that in reactor 78% more
cell mass was obtained compared to that of shake flask
experiment. In case of shake flask the growth rate is less
as the gas induction was due to only surface aeration. Ini-
tially sufficient dissolved oxygen would be available for
growth of organism as the medium was saturated with air.
The drop in dissolved oxygen was very rapid in the initial
phase and remained more or less same with the time. This
oxygen tension which decides the mass transfer coefficient
is responsible for the growth in the shake flask. Under these
conditions, the product yield was highest for 2 h of reac-
tion. In the case of bioreactor the KLa has been maintained
constant as it was obtained in shake flask due to the oxygen
tension. So the growth in the reactor was very high (78%
more) as there is no depletion of the dissolved oxygen. The
mixing in the bioreactor was more compared to shake flask
which maintains the sufficient oxygen levels. This could be
contributing for the formation of more cell mass. Aeration
and agitation must have provided a well mixed system which
may be equivalent to the optimum respiration rate required
for the organism for a better metabolic activity. In the case
of biotransformation even though organism is not growing
it requires oxygen to sustain itself as well as to perform
biotransformation activity. In the case of bioreactor more
yield of l-PAC and less yield of by-products were obtained
compared to that of shake flask experiment (Table 1). This
can be attributed to the high cell density, better distribution
of the cell mass over the reactor and higher optimal mass
transfer rates and to the fact that total cell mass present in
the bioconversion medium determines l-PAC production
[10]. In the case of bioreactor more cell mass was pro-
duced and it was subsequently used for giving more yield
of l-PAC. This highlights the need for the proper scale-up
criteria and need for proper pilot/bench scale experiments
coupled with the hydrodynamic requirement (mixing and
suspension) of the scaled up process. Scaling up directly
from the shake experiments could have led to a consider-
able over design for specific production rate which has been
avoided.
3.8. Conclusions
1. The scale-up of the l-PAC production from 100 ml
shake flask to 5 l fermenter has been done with KLa as
scale-up criteria.
2. Correlations for the mass transfer coefficients were
developed for both the growth and biotransforma-
tion media. The respiration rates in shake flasks were
used to estimate the mass transfer coefficient in the
fermenter. These mass transfer rates were used in
fixing the operating parameters like impeller speed
and air flow rates. This led to high production rates
in a fermenter as compared to shake flask indicating
the role of hydrodynamic parameters such as mixing
and suspension on the efficiency of the fermenta-
tion.
3. An optimum reaction time exists for the production of
l-PAC in both the shake flask and 5 l bioreactor which
is around 120 min and more product formation was
observed in the bioreactor.
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