a microwell platform for the scale-up of a multistep
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A microwell platform for the scale-up of a multistepbioconversion to bench scale reactors: sitosterol
side-chain cleavageMarco P Marques, Joaquim M. S. Cabral, Pedro Fernandes
To cite this version:Marco P Marques, Joaquim M. S. Cabral, Pedro Fernandes. A microwell platform for the scale-up of amultistep bioconversion to bench scale reactors: sitosterol side-chain cleavage. Biotechnology Journal,Wiley-VCH Verlag, 2010, 5 (4), pp.402. �10.1002/biot.200900098�. �hal-00552332�
For Peer Review
A microwell platform for the scale-up of a multistep bioconversion to bench scale reactors: sitosterol side-chain
cleavage
Journal: Biotechnology Journal
Manuscript ID: biot.200900098.R4
Wiley - Manuscript type: Research Article
Date Submitted by the Author:
10-Feb-2010
Complete List of Authors: Marques, Marco; IBB-CEBQ-IST
Cabral, Joaquim M. S.; IBB-CEBQ-IST Portugal Fernandes, Pedro; IBB-CEBQ-IST Portugal
Primary Keywords: Biochemical Engineering
Secondary Keywords: Biotransformation
Keywords: Microwell plates, Mycobacterium sp., scale-up
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Research Article ((6860 words))
A microwell platform for the scale-up of a multistep bioconversion to
bench scale reactors: sitosterol side-chain cleavage
Marco P.C. Marques, Joaquim M.S. Cabral, Pedro Fernandes*
1IBB-Institute for Biotechnology and Bioengineering, Centre for Biological and
Chemical Engineering, Instituto Superior Técnico, Av. Rovisco Pais, 1049-001 Lisboa,
Portugal
* corresponding author.
E-mail: [email protected]
Fax number: +351 218419062
Phone number: +351 218419065
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Abstract
The microwell-scale approach is widely used for screening purposes and one-pot
biotransformations, but it has seldom been applied to complex whole cell multistep
bioconversions, requiring prolonged incubation periods. The present study aims to
contribute to fill in this gap. The side chain cleavage of sitosterol to androstenedione
(AD) with Mycobacterium sp. NRRL B-3805 cells was used as model system, and focus
was given to the screening of suitable bioconversion media with 24-well microwell
plates. Results show that to perform this particular bioconversion growing cells are
preferred over resting cells due to higher conversion yields obtained in aqueous medium.
The use of resting cells may nevertheless present an interesting approach provided
catalytic activity is retained throughout successive runs. Maintaining suitable aeration
levels (air flow of 1ml.min-1
) allowed minimizing the decay of catalytic activity typically
observed alongside consecutive bioconversion runs with resting cells. Microwell plates
with dedicated oxygen and pH monitoring capabilities proved effective in media
development for complex multistep bioconversions using relatively slow-growing
bacteria. Under constant kLa (0.044 s-1
) similar AD production and dissolved oxygen
profiles were observed in microwell plates and in bench-scale reactor. Selection of a
suitable kLa value proved critical, since under lower kLa values scale-up proved
unsuccessful. The same pattern was observed when other scale-up criteria were evaluated
to perform the scale-up of this particular bioconversion. Results gathered seem to validate
the proposed approach “from microwell plate to bench-scale fermenter”.
Deleted: ¶
Deleted: to minimize
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Keywords: Biochemical Engineering, Mycobacterium sp., scale-up, bioconversion, kLa
1 Introduction
The use of small scale vessels (volumes below 100 ml) in order to speed up the
development fermentation/bioconversion is becoming widespread [1,2]. Shaken flasks,
test tubes and microwell plates (MWP) are extensively used due to the ease of
parallelization, and in the particular case of MWP, to the possibility of automation. They
have been shown to provide suitable platforms for the screening of biocatalysts, for the
collection of kinetic data, and in the early stages of process optimization (e.g. medium
development) [3-5]. These platforms have been further improved and their range of
application has increased, with the development of non-invasive optical fluorescence and
light scattering sensors for pH, dissolved oxygen, NADH and biomass monitoring [6-8],
and of devices for the control of pH and dissolved oxygen [9,10]. These technological
developments were also timely introduced to contribute to gaining insight on the
identification of the key parameters required to allow for reproducibility between shaken
vessels and bench-scale stirred reactors, a feature that only in a relatively recent period
has been suitably addressed [11-14]. There are hence few examples of scaling
fermentation or bioconversion processes directly from MWP to bench-scale bioreactor
partially due to the lack of instrumentation in traditional MWP [5,14-16].
Scale-up criteria are essential in bioprocess design given that the data generated are
translated rapidly and reproducible at larger scales. Most of the fermentation scale-up
techniques rely on maintaining constant a given parameter throughout the different
Deleted: Microwell plates
Deleted: hasten
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scales, viz. the volumetric power input, the impeller tip speed, the Reynolds number, the
mixing time or the volumetric oxygen mass transfer [17,18]. The volumetric oxygen mass
transfer is often the preferred approach when dealing with fermentative systems
involving aerobic microorganisms, and where oxygen supply may be a limiting factor
[19-22]. Islam et al. [14] and Micheletti et al. [21] used a constant volumetric oxygen
mass transfer for scale translation from MWP to bench scale/pilot bioreactor, of
processes for recombinant protein expression with the fast-growing E. coli and antibody
production in suspension cultures of VPM8 hybridoma cells.
The present work seeks to provide some contribution on these matters. An exploratory
and integrated perspective of the potential of MWP for bioprocess development, from
media selection to scale-up to bench bioreactor, according to suitable criteria, is provided.
The model system used was the selective cleavage of the side chain of β-sitosterol
performed by Mycobacterium sp. NRRL B-3805 cells. This is a well-established multi-
enzymatic process involving the use of eleven catabolic enzymes in a 14-step metabolic
pathway [23]. The product of this selective cleavage is 4-androstene-3,17-dione (AD),
which is a key intermediate in the production of pharmaceutical steroids. Previous studies
showed the validity of the use of MWP for performing this particular bioconversion with
free resting cells [24] and growing cells [25] of Mycobacterium sp. NRRL B-3805. A
comparison between these two systems is provided, highlighting the pros and cons of
their use.
2 Materials and methods
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2.1 Materials
Yeast extract and potato dextrose agar were obtained from Difco (Detroit, MI, USA).
Ammonium chloride was supplied by Merck-Schuchardt (Hohenbrunn, Germany),
glycerol (p.a. grade) was from Riedel-de-Häen (Seelze, Germany), β-sitosterol was from
Acros (Geel, Belgium) and Tween 20, 4-androstene-3,17-dione (AD), 1,4-androstadiene-
3,17-dione (ADD) and progesterone were obtained from Sigma (St Louis, MO, USA).
The SensorDish®
Reader, HydroDish®
and OxoDish®
MWP were from PreSens GmbH
(Regensburg, Germany). HydroDish®
plates presented a measurement range of pH 6 - 8.5
with a resolution of ± 0.05 and accuracy of ± 0.1 at pH=7. OxoDish®
had a measurement
range of 0-250% air saturation with a resolution of ± 2% and accuracy of ± 5% at 100%
air saturation. All other chemicals were of analytical or high-performance liquid
chromatography (HPLC) grade and purchased from various suppliers.
2.2 Microorganism and AD production conditions
Mycobacterium sp. NRRL B-3805 cells were maintained in potato dextrose agar slants
(42 g.L−1
). The inoculum was prepared as described by Staebler et al. [26]. Briefly, the
inoculum was pre-cultivated in complex medium, contained in 100 mL Erlenmeyer
flasks, with a headspace of 80%, under 200 rpm orbital shaking. A given volume of the
inoculum, corresponding to 10% (v/v) of the final volume, was transferred to the
production medium, once an optical density of roughly 0.9 (640 nm) was achieved. Trials
in shaken systems were performed in triplicates, at least, and were carried out in: i) 24-
well HydroDish®
(for pH monitoring) and OxoDish®
(for dissolved oxygen monitoring)
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MWP, where each well was filled with a total of 0.5 mL of production media, and MWP
were sealed with tapes (Excel Scientific, CA, USA), unless stated otherwise. ii) 250 mL
Erlenmeyer flasks filled with 25 mL medium. Production medium (used for simultaneous
cell growth and sitosterol bioconversion) was composed of either complex or defined
medium. Complex medium was composed of (g.L-1
) glycerol (10.0) yeast extract (10.0),
Tween®
20 (0.8), MgSO4.7H2O (0.2). Defined medium was composed of (gL-1
) glycerol
(20.0), NH4Cl (4.0), Tween®
20 (0. 8), MgSO4.7H2O (0.2), unless stated otherwise. Both
media were prepared in pH 7 sodium-potassium phosphate buffer (0.1 M) with initial
micronized β-sitosterol concentration of 2.4 mM (except for the pre-cultivation of the
inoculum, where β-sitosterol was absent from the medium). Trials were performed in
triplicates, at least, at 30ºC in orbital shakers Aralab Agitorb 2001C (Portugal) with 25
mm shaking diameter. When cell reuse was evaluated, after each bioconversion run of 24
hour, the cells were harvested and rinsed with buffer to remove traces of steroids. Fresh
medium was added and a new bioconversion run of 24 hour was carried out in the pre-
established conditions. Samples were taken every 12 hours and extracted with a two-fold
volume of a solution of progesterone (0.2 g.L−1
, internal standard) in n-heptane, for the
off-line quantification of biomass, glycerol (both in the aqueous phase), sitosterol and
AD(D) (in the organic phase [26]). In the MWP a sacrificial well approach was used.
Bench-scale fermentor batch trials were carried out in a 5-L bioreactor (Biostat B, B.
Braun, Germany) equipped with two Rushton turbines. Cell growth and bioconversion
were performed under the following general conditions: medium volume 4L, aeration rate
1 vvm, stirring speed of 220 and 470 rpm (resulting in kLa of 0.010 and 0.044 s-1
,
respectively). Temperature, initial pH, medium composition and inoculation volume
Deleted: w
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(10% (v/v)), were similar to the shaken systems. No significant changes were observed
for pH values throughout the whole incubation periods. Reactor dimensions are given in
Figure 1.
2.3 Oxygen mass transfer coefficients and volumetric power consumption
Oxygen mass transfer coefficients (kLa) were obtained by an enzymatic method in
accordance with Duetz and Witholt [27], for microwell plates and shaken flasks, and by
the gassing-out method applying the sulfite method according to John et al. [28], both for
microwell plates and bench-scale reactor .
In microwell plates, the oxygen mass transfer coefficient was determined for a fixed
shaking diameter (25 mm) at different shaking frequencies (0-300 rpm) and filling
volumes (250 – 1000 µL). kLa in bench scale reactor was determined in ion free medium
(water) and in different water:buffer ratios, as well as with different distances between
the two Rushton turbines (6.7, 8.8 and 15.4 cm).
Volumetric power consumption in bench scale reactors was determined according to
Van’t Riet [29] and Schmidt [17].
2.4 Assessment of influence of oxygen on the catalytic activity
The influence of oxygen on catalytic activity of free resting cells was assessed in standard
24-well microwell plates with magnetic stirring through the use of different air flows
(Figure 2). In a CERTOMAT®
H incubator (B. Braun, Germany) with controlled
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temperature (30ºC) and humidity (above 70%), the inner wells of the standard microwell
plate were magnetically stirred (Ikamag®
Reo, Drenzhal Electronic, Germany). Air was
supplied using a standard air compressor with controlled flow. Prior to feeding to the
bioreactors, air was bubbled through water for humidification, hence reducing
evaporation losses of the bioconversion medium. Additionally, microwell plates were
covered with a sandwich cover (EnzyScreen BV, Netherlands). Evaporation in the
microwell plates was controlled by weight on an hourly basis, and the medium was
supplemented with distilled water when evaporation loss was higher than 10%.
2.5 Statistical analysis
Data were treated using statistical analysis software (SPSS 14.0). The statistical treatment
was obtained by one way ANOVA, which was used to detect differences among
variables. Statistical confidence was set at 95%.
2.6 Analytical methods
2.6.1 Steroid analysis
HPLC analysis (Lichrospher Si-60 column, 5 µm particle size, Merck, Germany) with
1ml.min-1
isocratic elution was performed to determine substrate and products
concentration, with UV detection at 220 nm and 254 nm, respectively. The mobile phase
was composed of n-heptane and ethanol (90:10, v/v).. In all cases, ADD amounts were
vestigial, and, when quantifiable, followed the trend of AD formed.
Deleted:
Deleted:
Deleted: 5
Deleted: 5
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2.6.2 Carbon source analysis
Glycerol concentration was determined off-line by HPLC (Hitachi LaChrom Elite) with a
Bio-Rad Aminex Fermentation Monitoring Column (150 x 7.8 mm) by refraction index
with 0.8 mL.min-1
isocratic elution. The mobile phase was composed of 50 mM H2SO4
and the column was kept at 65 ºC.
2.6.3 Cell concentration
The protein concentrations of the samples was determined from the water phase using the
BCATM
Protein Assay Kit (Pierce, USA), after protein extraction by heating in a 1 M
NaOH solution for 20 minutes at 100ºC [30]. Protein concentration was converted to dry
biomass through a previously established calibration curve using mycobacterial cells,
where those two parameters were correlated.
3 Results and discussion
3.1 Setting up a well defined medium for the production of catalytically active microbial
cells
The microbial side chain cleavage of phytosterols is typically performed using complex
media [31-34]. However, the use of a defined medium enables the evaluation of the effect
of each component, favours the control of medium formulation and enhances
reproducibility. Successful examples of this strategy for sitosterol side-chain cleavage
have been recently implemented, yet based in a non-conventional environment [34-36].
Deleted: 5
Deleted: 5
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Given the promising results previously obtained with whole cells grown in complex
media [26], 24-well MWP were therefore used as a test bed for an exploratory screening
of a defined medium that would closely reproduce the performance of the complex
medium. The influence of different concentrations of glycerol and ammonium chloride,
present in the production media, on AD production was evaluated. Variation of AD
concentration was observed using the different combinations of nutrients. The best
outcome was obtained with the combination of (g.L-1
) glycerol (20.0), NH4Cl (4.0),
producing roughly 1 mM of AD, corresponding to a bioconversion yield of 43% after a
72 hour incubation period (Table 1). The time course of AD production in the defined
environment closely matched the trend observed when complex media was used as
production medium, (Fig. 3). The statistical similarity was of 0.50 (p-value). Comparing
growth rates of Mycobacterium sp. cells in both media, complex and defined medium,
they were statistically similar (0.13 ± 0.02 and 0.11 ± 0.02, respectively).
3.2 Using growing cells or resting cells.
Biotransformation of β-sitosterol into AD is possible due to the catabolic metabolism of
Mycobacterium sp. NRRL B-3805, independently of the primary metabolism of the cell
[23]. Both growing and resting cells of Mycobacterium sp. NRRL B-3805 are thus able to
produce AD from sterols [24,37,38]. In the former case, the selection of the operational
conditions, and in particular those related to aeration and oxygen transfer, has to take also
into consideration the requirements of the microbial cell for growth, since mycobacteria
are strict aerobes. Nonetheless, it is also known that carbon dioxide is likely to stimulate
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growth of mycobacteria [39]. When resting cells are used, the operational conditions are
not likely to be so demanding since growth constraints are ruled out. The same is likely to
occur when the role of the carbon source is considered, the potential need for a carbon
source is likely to be related to maintenance requirements and co-factor regeneration.
Considering the catalytic activity of both resting and growing cells over a 72 hour run
(Figure 4), there is a slight decrease in the yield in growing cells and an increase with
resting cells. This is primarily due to the in vivo catalytic system (growing cells have a
metabolic burden that comprises cell growth and sterol bioconversion whereas in resting
cell the metabolic burden is restricted to sterol bioconversion metabolism and general cell
maintenance metabolism). Moreover the AD production is proportional to the number of
catalytic active cells. This proportion is observed till a cell concentration of roughly 30
gL-1
[26]. On the other hand, with resting cells, the biomass concentration is constant
along the biocatalytic run (25 g.L-1
).
Wang et al. [40] showed that under carefully selected environments, designed for either
case, close final ADD yields can be obtained from similar initial substrate concentration
in batch runs using either growing or resting cells. Nonetheless, the currently
implemented processes for large-scale biobased production of AD(D) from sterols rely
on the use of growing cells of Mycobacterium sp. [41, 42].
In order to allow the reuse of resting cells in consecutive runs, bioconversion was
performed under forced aeration (Figure 5), given the encouraging results obtained when
operating in non-conventional environments [38]. .
The results show that there is a strong correlation between AD production and the level
of aeration. Runs performed under forced aeration (air flow of 1 mL.min-1
) produced
Deleted: Tentatively, in order to improve the performance of resting cells
over a bioconversion run (with cell
reuse), different strategies were assessed,
namely forced aeration (Figure 5), given
the encouraging results obtained when
operating in non-conventional environments [38]
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roughly 50 % more AD than those under either non-forced aeration (medium aeration –
0.1 mL.min-1
) or non-aerated conditions (low aeration - 0.01 mL.min-1
). Curiously under
nitrogen atmosphere relatively significant amounts of AD were produced. This suggests a
basal level of catabolic activity of sterol-side chain pathway, despite the low oxidative
potential. It is likely that the pool of cofactors available in the cells and required for the
bioconversion to proceed [23] is used by the catalytic pathway until total depletion.
Promoting operational conditions that increasingly favor oxygen transfer into the
bioconversion media led to a concomitant increase in AD production, up to a 2-fold
increase at the end on bioconversion runs when an air flow of 1 mL.min-1
is added, as
compared with the basal level. Although forced aeration allowed for an increase in the
specific activity of resting cells, sustained reuse of resting cells did not prove viable and
the methodology for performing runs under aeration proved cumbersome and labor
intensive. Therefore, throughout further work, scale-up studies are performed using
growing Mycobacterium NRRL B-3805 cells.
3.3 Scale-up studies
The selected medium was further used in order to obtain proper scale-up criterion to
perform scale translation from microwell plate to bench-scale reactor. KLa was primarily
chosen as scaling criterion given the promising results observed in previous works
involving scaling-up applied to cell-growing microbial systems [14,21,25].
3.3.1 Determination of kLa in microwell plates
Deleted:
Deleted: .
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In order to characterize the oxygen mass transfer coefficient at one fixed shaking
diameter (25 mm) at different shaking frequencies (0-300 rpm) the chemical sulfite
method was applied [28] and compared with off-line data gathered by an enzymatic
method [27]. A filling volume of 500 µL was used and plates were equipped with a lid to
emulate operational conditions. The use of adhesive tape to seal MWP prevents medium
evaporation and cross-contamination due to spill-over. The use of the sealing tapes leads
to airtight closure of the wells decreasing oxygen transfer to the reactor well and
consequentially to the reaction medium [43]. The kLa was determined by a simple
stoichiometric balance to the oxidation reaction:
[ ] [ ] [ ]( ) OUROOakdt
OdL −−= ∗
222
(1)
where [O2]* represents the dissolved oxygen concentration at equilibrium and OUR is
the oxygen uptake rate of the chemical reaction. Data was obtained in accordance with
[28].
Previously, Marques et al. [25] showed that with increased filling volume there is a shift
in kLa, decreasing from 0.096 s-1
(250 µL) roughly to 0.028 s-1
(1000 µL) at 300 rpm. The
oxygen transfer depends on the gas–liquid interfacial area, which is characteristic of
surface-aerated reactors, being higher when the filling volume is lower. Moreover, a clear
difference was observed between unsealed and sealed microwell plates in kLa, in a
magnitude of 50% due to an decrease in oxygen transfer rate.
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The experiments were performed at 30ºC, with sealed MWP to emulate long term
incubation periods (data not shown).
The kLa value increased with increased shaking frequency reaching a maximum of 0.048
± 0.08 s-1
at 300 rpm. Values of oxygen mass transfer obtained off-line by an enzymatic
method [27] are in the range of the kLa obtained on-line with the oxygen monitored
microwell plates. Similar values for kLa were obtained by Marques et al. [25], Islam et al.
[14] and Kensy et al. [44] for 24-well microwell plates, who reported kLa values ranging
from 0.05 s-1
to 0.07 s-1
for the experimental conditions and microbial systems evaluated
[14,25,44].
Under the selected conditions for growing Mycobacterium sp. NRRL B-3805 cells in 250
mL unbaffled shaken flask (250 min-1
at 25 mm shaking diameter) the kLa obtained was
of 0.044 s-1
.
Scale-up trials were performed also at lower kLa values (non-optimized conditions),
namely 0.010 s-1
.. Productivity levels are expected to decrease since there is a deviation
from the previous established optimal conditions.
3.3.2 Determination of kLa in bench-scale reactor
In stirred reactors, volumetric power consumption and superficial gas velocity are major
correlation coefficients for kLa [29]. Therefore, equations of the following type are
frequently found in the literature:
( ) 3
2
1
C
s
C
L vV
PCak
= (3)
Deleted: .
Deleted: The kLa values used are in the range of 0.005 and 0.055 s-1 observed for similar systems by other authors
[13,28,45]
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where P = gassed power input, V = liquid volume, vs = gas superficial velocity and C1, C2
and C3 are constants. The oxygen mass transfer coefficient can be influenced by vessel
geometry parameters [29]. Despite the existence of several correlations, an exact value of
kLa was sought for this particular reactor assembly. Since the reactor used in this study
has a dual Ruston turbine system, the influence of the spacing between the turbines was
assessed on kLa in order to find the ideal position to maximize kLa and mixing time
(homogenization dynamics) [45] while minimizing foam formation (decrease in oxygen
availability; cell migration, due to their hydrophobic nature, towards foam as well as cell
immobilization onto excess antifoam).
Analysing kLa data, independent of the position of the upper RT, the kLa increases with
increased stirring speed. Moreover, there is not a clear difference in terms of kLa values
between turbine positions, confirmed by the p-value of 0.904. Nonetheless, with the
upper turbine placed at 15.4 cm (maximum spacing tested) foam formation was
enhanced. This spacing was therefore left out in further work, where a spacing of 8.8 cm
was used instead.
In the overall, statistical comparison of the correlations tested with experimental data
produced adequate matches, with some minor exceptions possibly due to different ionic
medium strength and geometrical parameters of the reactors. By raising the ion
concentration in solution, kLa increases [29]. The influence of the ionic concentration was
tested in buffered systems up to a concentration of 0.1 M phosphate buffer (Figure 6).
Oxygen mass transfer coefficient in the production medium was determined being in the
range of 10 – 15% of the 0.1 M phosphate buffer.
Deleted: 46
Deleted: until
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Oxygen mass transfer coefficient increases with the ionic strength of the solution for the
same stirring speed. For buffer concentration used in this particular study, there is a
correlation of 0.78 (p-value) between experimental data and literature model [29]. This
indicates that Van’t Riet correlation can describe kLa in culture medium. Nonetheless,
scale-up based on the mass transfer rate in gas-liquid stirred tanks is challenging since
larger scale tanks can suffer from zones of oxygen depletion, particularly where there is
an oxygen sink, e.g. through biochemical reactions. It is impossible to keep all quantities
constant at different scales, but it is feasible to maintain a couple of variables i.e. a
combination of Pg/V and either Flg (aeration number), vvm (air flow) or vg (gas
superficial velocity). It is generally accepted that constant Pg/V should be maintained,
since it directly affects the local energy dissipation rate, which is the key hydrodynamic
variable in the breakage and coalescence kernels [46]. Therefore, kLa predicted using
equation (3), which suggests that a dependence on Pg/V and vg, may not stand for
bioreactors of a size different to the one where the correlation was produced [46].
3.3.3 Scale- up based on kLa
The bioconversion system was scaled-up from 24-well microtiter plate to 5 L bench-scale
in a 6000 scale-up fold base on maintaining kLa. The conditions were set at optimized
bioconversion conditions in shaken systems in aqueous medium [24] corresponding to
0.044 s-1
and at a lower kLa, namely 0.010 s-1
, corresponding to non-optimized
bioconversion conditions.
When analyzing the AD production resulting from the scale-up (Figure 7) there is clearly
a difference between the two set of conditions. At the highest kLa values tested specific
Deleted: 47
Deleted: 47
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AD production is statistically identical in all scales (p-value of 0.90). At lower values, not
only specific AD production is lower as expected, since non-optimized conditions were
used, but data for AD production are also statistically unrelated when the different reactor
systems are considered. In the recent papers by Islam et al. [14] and by Micheletti et al.
[21] the volumetric oxygen mass transfer was effectively used for scale translation from
MWP to bench scale/pilot reactor, of processes for recombinant protein expression with
the fast-growing E. coli and antibody production in suspension cultures of VPM8
hybridoma cells. The results indicate that kLa is a suitable criterion for scaling-up the
bioconversion of β-sitosterol to AD.
3.3.4 Time course of typical bioconversion runs in MWP and in bench fermenter
Using the kLa similarity as criterion for scale up, bioconversion runs were performed in
pH and oxygen monitored MWP and in a 5 L stirred reactor (Fig. 8). Online data for
oxygen depletion, determined by partial oxygen pressure, pO2, were collected in MWP
and compared with data gathered in 5L bench scale reactor. As for carbon source
depletion, AD formation and biomass production were assessed off-line in order to have
comparable measurements with similar error.
Similar studies were performed by Marques et al. [25] with this particular bioconversion,
but using a complex bioconversion media, and only a single kLa value as scaling
criterion. In the present work, where a defined medium was used and operation was
carried out under a lower kLa, product yield and productivity roughly matched the results
previously obtained with the complex medium..
Deleted: likely the most
Deleted: criteria
Deleted: purposes of
Deleted: .
Deleted: Nonetheless, results obtained were preliminary and did not take into
account optimized bioconversion medium
or conditions
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In the overall, the time course of the bioconversion at 0.044 s-1
follows a relatively
similar trend in both MWP and bench bioreactor, but monitored parameters do not fully
overlap. There is an initial sharp decay of oxygen in the bench reactor, which was not
observed in the MWP system. The difference is possibly due to the forced aeration and to
enhanced mixing conditions. Due to their marked hydrophobic nature, Mycobacterium
sp. cells tend to form clusters in aqueous systems. This influences mass transfer of
oxygen and substrate towards cells. In MWP, due to reduced volume, the influence of
these factors is more pronounced. Statistically, oxygen profiles are not equal (p value of
7.04×10-31
).
When the remaining monitored parameters are considered, again despite of a similar
trend being observed, a gap between values obtained in the MWP system and bench-scale
reactor was observed. Once again, intrinsic hydrodynamics conditions to the reactors can
be accounted for this behaviour. Even so, p-value scores for biomass, AD production and
carbon source depletion are 0.18, 0.15 and 0.33, respectively.
Maintaining kLa value of 0.01 s-1
constant no overlapping or similar trend of the time
course of bioconversion was observed. The profile for AD production, carbon source
depletion and cell growth (protein profile) were not matched in MWP and in bench
bioreactor. Moreover, the highest values obtained for these profiles were lower than the
ones obtained at higher kLa.
3.3.5 Other parameters
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The volumetric power consumption and Reynolds number were also calculated (Table 2).
Under the operational conditions required for the established oxygen transfer rates,
deviations in the volumetric power input and Reynolds number can be observed. Despite
recent progresses [49,50], the knowledge of the basic fluid mechanics in shaking systems,
particularly at the MWP scale, still lags behind the knowledge gathered for stirred
systems and, in the overall, for large scale systems. If we consider traditional Reynolds
transitional states, two distinct “regimes” can be observed. One corresponding to laminar
flow regime reached in MWP and the other corresponding to turbulent flow regime
obtained in Erlenmeyer flasks and in the bench scale reactor (a 20 fold increase is
observed in terms of Reynolds number). For scale-up purposes, in this particular
bioconversion system, Reynolds number can not be used since, in order to achieve the
same values in microwell plates, an increment in shaking frequency must be performed.
However, at higher shaking frequencies (300 rpm or higher), maintaining the shaking
amplitude (25 mm), splashing phenomena occurs. An alternative to overcome this
situation lies in the use of baffles inside the MWP wells. Between the two conditions
tested no significant difference is observed in terms of flow regime. Moreover, at larger
scales, power requirements to obtain similar Reynolds numbers are high and in terms of
overall process economics avoided. In industry, there are increased restrains on
environmental regulations, designing process that enables low impacts on environmental
issues (including dispending less amount of energy). On the other hand, high shear stress
could lead to lower production yields due to stress-induced cell damage [51].
Regarding volumetric power input there is also a divergence between the different scales
tested. Higher values are reached by the bench scale reactor due to the higher power
Deleted: 50
Deleted: 51
Deleted: higher
Deleted: 52
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necessities to stir at higher frequencies. Actually the use of volumetric power input as
scaling criterion is somehow limited, since fermentations requiring high energy inputs,
result in high shear stress in the larger scale stirred vessels and high costs are associated.
Besides, information is rather scarce when microwell plate scale is considered.
Accordingly, data was not available for microwell plate in the conditions tested.
Nonetheless, and despite no direct comparison can be made, Zhang et al. [49] established
a range of 0.07 to 0.1 kW.m-3
for 24-well microwell plate at 500 to 1500 rpm respectively
for shaking amplitude of 3 mm. At 1000 rpm the power consumption measured was 72
W.m-3
.
4 Concluding remarks
MWP were effectively used for the selection of a synthetic bioconversion medium for
sitosterol side-chain cleavage with growing cells able to closely replicate results obtained
with a complex media. A 6000-fold scale-up, from MWP to bench scale reactor, was
effectively performed for this bioconversion, based on kLa similarity, although such
reproducibility is apparently limited to given values of this parameter. This feature may
be related to highly differentiate mixing and mass transfer patterns when the two systems
are operated under conditions corresponding to relatively low kLa values. Under the
selected conditions for scale-up, the time course of typical bioconversion runs in either of
the scales displayed roughly similar trends when biomass production, relative product
yield and substrate depletion were considered.
Deleted: Data
Deleted: 50
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In an economical view, this allows not only cost reduction with associated equipment,
reagents and handling but also it speeds up the development and optimization of
bioconversion processes.
Growing cells displayed a higher specific catalytic activity than resting cells. The use of
the latter in consecutive runs was evaluated, and despite some preliminary encouraging
results, further studies are clearly required, since the feasibility of this approach is far
from satisfactory.
Oxygen mass transfer coefficient is a suitable scaling-up criterion for bioconversions
using Mycobacterium sp. NRRL B-3805. Maintaining a value of 0.044 s-1
is was possible
to uphold productivity yields over a 6000 fold scale-up.
These results show that it is possible to use specific microwell plates, with online
measurement of oxygen complex whole cells bioconversion processes, to mimic runs
performed at bench-scale.
M.P.C. Marques and P. Fernandes acknowledge Fundação para a Ciência e Tecnologia
(Portugal) for financial support in the form of PhD grant SFRH/BD/24433/2005 and
program Ciência 2007, respectively. This work was partially funded by research project
PPCDT/SAU-MMO/59370/2004 from Fundação para a Ciência e Tecnologia (Portugal).
Ricardo Pereira is acknowledged for his help on the reactor maintenance.
The authors have declared no conflict of interest.
Formatted: Font: Italic
Formatted: Font: Italic
Deleted: hasten
Deleted:
Deleted: themselves
Deleted: the most
Deleted: criteria
Deleted: Acknowledgements¶
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Caption to figures
Figure 1. Reactor dimensions. A – 5 l bench scale reactor; B – 250 ml shaken flasks and
C – 24-well microtiter plate. Dimensions: Ab=2.7 cm; Af=8.1cm; Wz=3.0cm; Di=7.7 cm;
Li=2.0 cm; Wi=1.5 cm; Wb=1.6 cm; Wk=4.6 cm.
Figure 2. Experimental set-up for the analysis oxygen influence on catalytic activity of
resting cells. A - Single well of 24-well microtiter plate with magnetic stirring and forced
aeration. B - Experimental set-up - a - air compressor, b – air humidifier, c - air
distributor, d - microtiter plate, e - magnetically stirrer plate and f - incubator.
Figure 3. Bioconversion progress for AD production using Mycobacterium sp. NRRL B-
3805 cells growing in defined (triangles) or in complex (squares) medium, incubated in
MWP filled with 500 µL reaction volume, under 250 rpm orbital shaking.
Figure 4. Catalytic activity of growing (�) and resting (�) cells along a 72 hour run.
Protein content: Growing cells – 3.20 g.L-1
(24 ), 8.74 g.L-1
(48 ) and 11 g.L-1
(72 ).
Figure 5. Influence of oxygen on the catalytic activity of resting cell along several cell
reuses. High aeration –1 ml min-1
aeration; Medium aeration - 0.1 ml min-1
aeration
(mimicking cotton plug); Low aeration – 0.01 ml min-1
and No aeration – anoxic
headspace. Each bioconversion run lasted for 24 hours.
Deleted: ,
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Figure 6. Oxygen mass transfer coefficient in dual Rusthon Turbine stirred tank (RT
distance of 8.8 cm) at different ionic strength medium at 30ºC, with kLa shown as
function of the stirrer speed. Growth medium corresponds to buffer 0.1 M.
Figure 7. AD production in MWP, shaken Erlenmeyer flask and batch stirred tank
reactor, BSTR, using kLa as scale-up criterion. kLa values: 0.044 s-1
(white bars) and
0.010 s-1
(grey bars).
Figure 8. Online monitoring of Mycobacterium sp. NRRL B-3805 growth in 5L bench
scale reactor (solid line) and in 24-well MWP (dashed line) at 0.010 s-1
(A, B and C) and
0.044 s-1
(D, E and F), Off-line monitoring of biomass (A - D), AD production (B - E)
and carbon source consumption (C - F). � - Bench scale reactor; △ – Microwell plate.
Deleted: 6.7
Deleted: the
Deleted: various reactor and/or agitation configuration
Deleted: a
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Table 1 - Example of screening of suitable composition of defined media for AD
production from sitosterol using Mycobacterium sp. NRRL B-3805 cells at 30ºC. Data
given report on the effect of glycerol and ammonium chloride concentrations in AD
production (mM). Standard deviation did not exceed 7%.
Table 2 – Summary of parameters obtained for runs performed at the microwell, shaken
flask and laboratory scales.
kLa 0.010 s-1 kLa 0.044 s-1
Reactor type
P/V
(W.m-3
)
Re
P/V
(W.m-3
)
Re
24 MTP 0.5 ml Nd 847 nd 1411 (2)
250 ml flasks 50 ml 280 (3)
20397 330 (1)
33994 (2)
5L reactor 4 L 1843 60210 2620 67706
nd = not determined; (1) – obtained by Büchs et al. [48] at the same conditions; (2) – obtained by applying
the correlations of Lotter and Büchs [49]; (3) – obtained by Zhang et al. [50] at the same conditions.
NH4Cl (gL-1)
2 4 8 12
5 0.75 0.78 0.86 0.86
10 0.82 0.84 0.89 0.86
20 1.04 1.05 1.00 0.97
Gly
cero
l (g
.L-1
)
40 0.87 0.88 0.91 0.97
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Biotechnology Journal
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Biotechnology Journal
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Biotechnology Journal
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Biotechnology Journal
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For Peer Review
190x254mm (96 x 96 DPI)
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Wiley-VCH
Biotechnology Journal
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Wiley-VCH
Biotechnology Journal
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