miniature bioreactors for automated high-throughput bioprocess design

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Biotechnol. Appl. Biochem. (2005) 42, 227235 (Printed in Great Britain) doi:10.1042/BA20040197 227

Miniature bioreactors for automated high-throughputbioprocess design (HTBD): reproducibility of parallel fed-batchcultivations with Escherichia coli

Robert Puskeiler*1, Andreas Kusterer*, Gernot T. John and Dirk Weuster-Botz*

*Lehrstuhl f ur Bioverfahrenstechnik, Technische Universitat M unchen, Boltzmannstrasse 15, 85748 Garching, Germany, andPreSens GmbH, Regensburg, Germany

To verify the reproducibility of cultivations of Escheri-

chia coli in novel millilitre-scale bioreactors, fully auto-

mated fed-batch cultivation was performed in seven

parallel-operated ml-scale bioreactors with an initial

volume of 10 ml/reactor. The process was automati-cally controlled by a liquid-handling system responsible

for glucose feeding, titration and sampling. Atline ana-

lysis (carried out externally of the reaction vessel with a

short time delay) comprised automated pH and atten-

uance measurements. The partial pressure of oxygen

(pO2) was measured online by a novel fluorimetric sen-

sor block measuring the fluorescence lifetime of fluoro-

phors immobilized inside the millilitre-scale bio-

reactors. Within a process time of 14.6 h, the parallel

cultivation yielded a dry cell weight of 36.9+0.9 g l1.

Atline pH measurements were characterized by an

S.D. of


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228 R. Puskeiler and others

Table 1 Parameters used

Parameter Definition Units

cs Substrate concentration g l1

DI Impeller diameter mD Attenuance at wavelength Unitless

DCW Dry cell weight g l1

DO Diss olved oxyge n concentr ation Air satur ation (%)

kLa Ox yge n tr ansfer coefficient s1

N Impeller speed s1

NP Power number Unitless

M Mass flux substrate g h1 l1

OTR Oxygen transfer rate g h1 l1

OUR Oxygen uptake rate g h1 l1

P Power W

pO2 Par tial pressure oxyge n Air satur ation (%)

Liquid density kg m3

wavelength of 600 nm was found. pH can be monitored by

a solid-state pH-sensor chip mounted on the circuit board.The maximum reported attenuance, D, was 2.4 after a batch

cultivation without pH control.

A third approach is based on a magnetically driven gas-

inducing impeller mounted in a baffled bioreactor with a re-

action volume of 815 ml [13]. The design as a gas-inducing

impeller does not require gas sparging in the reaction vol-

ume, thus the individual control of gas inlet fluxes in the

range of ml min1 is not necessary. The maximum kLa was

characterized to be >0.4 s1. An automated fed-batch

cultivation in such a millilitre-scale bioreactor reached a

DCW of 20.5 g l1 in a mineral medium with glucose as the

sole carbon source. The growth ofEscherichia coliwas shown

to be equivalent to a reference cultivation in a laboratory-

scale stirred-tank bioreactor with a working volume of

3 litres [14].

The present study is the first to report the application

of the latter approach for the parallel cultivation of E. coli,

aiming at the verification of the reproducibility of parallel

cultivations. Therefore seven gas-inducing bioreactors at a

millilitre scale are operated in parallel in a reaction block

providing an electromagnetic drive and heat exchangers. The

automation of titration, feeding and sampling is realized by

a liquid-handling system. An integrated MTP reader enables

automated pH and attenuance monitoring. The process is

conducted with pH control in a fed-batch operation mode.Moreover, integrated fibre-optic sensors in the ml-scale bio-

reactors allow the parallel monitoring of pO2 [partial pres-

sure of oxygen (percentage air saturation)] with a frequency

of six data points/min per reactor. The parameters used in

this paper are defined in Table 1.

Materials and methods

The millilitre-scale bioreactors

The baffled bioreactors with a nominal volume of 815 ml

were produced from polystyrene by die casting (H+ P

Figure 1 Cross-section of a bioreactor mounted in the reaction block

The lower part of the reaction block is equipped with the magnetic drive and

heat exchangers to control the reaction temperature. The upper part of the

reaction block contains heat exchangers to enable cooling of the headspace of

the bioreactors thus limiting ev aporation. The sterile barrier provides the gas

inlet andthe gasoutlet, whichsimultaneouslyservesas sampling port.The shafts

on which the impellers rotate freely are screwed into the sterile barrier.

Labortechnik, Oberschleissheim, Germany). The gas-

inducing impellers were machined from polyether ether

ketone (PEEK) by a CNC (Computer Numerical Control)

milling machine (H+ P Labortechnik). The impellers are

equipped with samarium/cobalt (SmCo) permanent magnets

(Fehrenkemper Magnetsysteme, Lauenau, Germany). A

cross-section of one bioreactor mounted in the reaction

block is depicted in Figure 1. The dimensions of the bio-reactor, the shaft and the gas-inducing impeller are detailed

in Figure 2.

Reaction block

The reaction block (dimensions: lengthwidth height,

310 mm 240 mm 140 mm) fits a maximum of 48 bio-

reactors. The reaction block is equipped with the magnetic

drive, heat exchangers and a sterile barrier providing the

gas inlet and the gas outlet, which simultaneously serves

as sampling port. The shafts are screwed into the sterile

barrier. The gas-inducing impeller rotates freely on the shaft.

Maximum impeller speed can be set to 4000 rev./min, cor-responding to 66.6 Hz.

CFD (computational fluid dynamics) simulation of

hydrodynamics in the millilitre-scale bioreactor

The commercial CFD software package CFX-5.7 (Ansys,

Otterfing, Germany) was used for the modelling of the

hydrodynamics of the millilitre-scale bioreactor filled to a

liquid height of 40.5 mm corresponding to a liquid volume

of 11.2 ml. The code is based on the finite volume technique.

The reactor volume is resolved with a total number of

170000 hybrid grid control volumes. For each of these

C 2005 Portland Press Ltd


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Reproducibility in miniature bioreactors 229

Figure 2 Major dimensions of the ml-scale bioreactor, the shaft and the

gas-inducing impeller as they are mounted inside the reaction block

All the dimensions shown are in mm.

volumes, a full set of transport equations for mass andmomentum is solved in a steady-state model approach. To

represent the rotating parts of the reactor, the multiple

reference frame technique is used with a rotating inner

stirrer volume and a stationary outer vessel volume. Turbu-

lence effects in the fluid flow are prescribed using the SST

(shear-stress-transport) model. The simulation is performed

in a single-phase status using the fluid properties of water at

ambient temperature.

Organism and medium

The strain used for cultivation was E. coli K12 (DSM 498),

purchased from the German Collection of Micro-organismsand Cell Cultures (Braunschweig, Germany).

Seed cultures and bioreactor cultivation of E. coliwere

carried out in a defined mineral medium of the following

composition (adapted from [15]): glucose (15 g l1),

KH2PO4 (13.3 g l1), (NH4)2HPO4 (4.0 g l

1), MgSO4

7H2O (1.2 g l1), CoCl2 6H2O (2.5 mg l

1), MnCl2 4H2O

(15 mg l1), CuCl2 2H2O (1.5 mg l1), H3BO3 (3.0 mg l

1),

Na2MoO4 2H2O (2.5 mg l1) and zinc acetate dihydrate

(13 mg l1), citric acid monohydrate (1.86 g l1), EDTA

(8.4 mg l1), ferric citrate (12.5 mg l1) and thiamine hydro-

chloride (4.5 mg l1). As the antifoam agent, CLEROL 265

(0.1 ml l1; Cognis, Dusseldorf, Germany) was used.

Feeding solution for fed-batch cultivation contained

glucose (475 g l1), MgSO4 7H2O (20 g l1), EDTA

(13.0 mg l1), CoCl2 6H2O (4.15 mg l1), MnCl2

4H2O (24.9 mg l1), CuCl2 2H2O (2.5 mg l1), H3BO3(5.0 mg l1), Na2MoO4 2H2O (4.15 mg1 l

1), zinc acetate

dihydrate (21.7 mg l1) and ferric citrate (47.6 mg l1).

Seed culture

The 50 ml seed culture in a 500 ml unbaffled shake flask was

inoculated with 0.067% of a frozen E. coli cell stock and

incubated at 37C with shaking at 200 min1 for 16 h (Multi-

tron; Infors, Bottmingen, Switzerland) until a DCUV650 of 3.0

(Genesys20; Thermo Electron, Dreieich, Germany) was

reached.

Millilitre-scale bioreactor cultureThe millilitre-scale bioreactors were inoculated with 100 %

seed culture corresponding to an initial DCUV650 (attenuance

in the cuvette photometer at 650 nm) of 3.0. The initial

reaction volume was 10 ml. Stirrer speed was initially set

to 2000 min1. The gas flow through the sterile gas cover

consisted of oxygen-enriched air. The volumetric fluxes of air

and oxygen were controlled by thermal mass flow control-

lers (Brooks Instruments, Veenendaal, The Netherlands).

Total gas flow was initially set to 4.8 litres min1 air cor-

responding to 0.1 litre/min per reactor. During batch cul-

tivation, the stirrer speed was manually increased to main-

tain pO2 >20 % air saturation. During fed-batch culture, the

volumetric fraction of oxygen was additionally increased up

to a total oxygen concentration of 50% in the inlet gas

flow. pH was automatically controlled at 6.8 by intermittent

addition of 12.5 % (v/v) NH4OH.

Pipetting precision

The pipetting precision of the liquid handler was verified

gravimetrically for the pipetting of feed solution, with a

volume of 18.7 l corresponding to the highest feed flow of

13.3 g h1 l1. The density of the feed solution was gravi-

metrically determined to be 1.19 g cm3. The test com-

prised ten pipetting steps per pipette tip. The deviation of

the tips from the desired pipetting volume was taken intoaccount for the subsequent pipetting test.

Monitoring of cell growth

Cell growth was automatically monitored atline in MTPs by

measuring the attenuance, DMTP650 (microtitre-plate attenu-

ance at 650 nm), at a wavelength of 650 nm. Samples of

10 l were automatically diluted to 1:20 with PBS (8 g l1

NaCl, 0.2 g l1 KCl, 1.44 g l1 Na2HPO4 and 0.24 g

l1 KH2PO4, pH 7.4) by the internal pumps of the MTP

reader. DCW was estimated from DMTP650 with a second-

order polynomial function allowing a measuring range of



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undiluted samples between 0.16 and 1.74 DMTP650 units with

a mean CV (coefficient of variance), R2, of 0.995. Taking the

dilution factor into account, the resulting DCW measuring

range is 3.248.0 g l1.

Reference attenuance measurements (DCUV650) werecarried out up to a process time of 8.77 h, i.e. the pre-

sence of the experimenter. A single beam photometer

(Genesys20) with plastic cuvettes of 1 ml volume was used.

Measurements were carried out at a wavelength of 650 nm.

Monitoring of pH

Commercially available MTPs containing pH-sensitive sensor

spots (Hydroplate; Presens, Regensburg, Germany) were ap-

plied for automated atline pH measurements in an MTP

reader (Fluostar Galaxy; bmg labtech, Offenburg, Germany).

The hydroplates were calibrated with cultivation medium

adjusted to six different pH values covering the measuringrange between pH 4 and 9.

Monitoring of DO

DO was monitored online with fluorimetric sensor spots

(Presens) immobilized on to the reactor bottom with the sili-

cone rubber compound 692542 (RS Components, Corby,

U.K.). The fluorescence decay time of the fluorophors was

measured with a prototype fluorescence reader (Presens)

containing eight separate light sources and photodiodes

integrated into a sensor block fitting beneath the reaction

block. Measurements were performed with a frequency of

six data points/min per reactor.

Process control

All sampling and feeding steps as well as the MTP move-

ments were programmed with the liquid-handler control

software (Gemini; Tecan, Crailsheim, Germany). Sampling

of 50 l h1 was carried out by a liquid-handling system

(RSP 150; Tecan) with a frequency of 1 h1 during batch

growth. An additional sample was taken 30 min after the

initial glucose was metabolized. During fed-batch growth,

samples were taken each time the mass flow of feed sol-

ution was increased. From the process time of 6.9 h until

the end of the cultivation, samples were taken with a fre-

quency of 2 h

1. The samples were pipetted into an MTPand then transported to the plate reader (Fluostar Galaxy)

that measured pH in undiluted samples and attenuance

DMTP650 in 1:20 diluted samples. Communication with the

MTP reader was programmed in LabView (version 6.1;

National Instruments, Munich, Germany). LabView VIs

(virtual instruments) were started from the command line

in the liquid-handler control software. Closed loop control

for pH was achieved by extracting pH values from the MTP

reader data files, processing them in LabView and creating

pipetting instructions for the liquid-handler control software

that were regularly loaded into the control program. pH

was controlled at 6.8 with a simple proportional controller

calculating the required amount of base for any single reactor

by applying a titration curve that was created in advance.

No addition of acid was implemented into the pH-control

algorithm. During fed-batch growth, the titration agent wasadded with a frequency of 3 h1. After each sampling step, the

pH controller increased the base mass flow proportionally

according to the measured pH deviation from the set

point. After atline measurements, the diluted samples were

recovered for manual attenuance measurements in a single

beam cuvette photometer (DCUV650). The MTP was cleaned

in an MTP washer and re-used to receive the subsequent

samples. After 8 h of process time, a new hydroplate was

used to maintain high reproducibility of pH measurements.

After the consumption of the initial glucose at a pro-

cess time of 4.9 h, feeding was started intermittently with a

frequency of 15 h1

corresponding to a feeding interval of4 min. After 40 and 90 min, the initial feed flow of 4.0 g

h1 l1 was increased to 7.5 and 13.3 g l1 h1 respect-

ively. Hence, the highest feeding volume was 18.7 l for

each pipetting step.

Results and discussion

Hydrodynamics of the millilitre-scale bioreactor

Figure 3 shows the results of the single-phase CFD simul-

ations at an impeller speed of 2800 rev./min. Owing to the

rotational movement, the liquid phase is accelerated along

the diagonal outward pumping channels, thereby reaching a

top speed of>2.0 m s1. The strong radial flow induced

by the impeller leads to a circulating flow field above and

below the impeller comparable with the flow field of a single

Rushton turbine. As a result of the acceleration of the liquid

phase in the diagonal channels, a constant negative pressure

builds up in the vertical bottom channel of the impeller as

shown in the spatial distribution of the simulated pressure in

the liquid phase. The peak negative pressure (2000 Pa) is localized in the zone where the

outward pumped liquid encounters the baffles. The zones

of highest positive or negative pressure correspond to the

zones of highest energy dissipation. Since the integral of

the simulated energy dissipation over the liquid volume

generally underestimates the power input [16], this variable

was calculated by solving the momentum balance of the flow

field. This approach yields a mean power input of 21.9 W l1

for the millilitre-scale bioreactor at 2800 min1. The single-

phase power number NP = 3.7 is obtained by assuming a



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Figure 3 Results of single-phase CFD simulations of axially projected liquid speed and pressure in the millilitre-scale bioreactor for an impeller speed of

2800 rev./min

The liquid speed is represented in a vertical cross-section through the centre of the bioreactor. The pressure distribution is represented in a vertical cross-section

lying 0.7 mm in front of the centre of the bioreactor in order to display the difference between positive pressure in front of the baffles (right) and negative pressure

behind the baffles (left).

liquid density of= 998 kg m3 in the equation:

Np = P/ N3 D

5

I

Parallel fed-batch cultivation

After three optimization cycles, the accuracy of the eight

pipette tips of the liquid-handler was increased from the

initial value of 87.3 to 99.9 % with a precision of 3.0 % of 10

subsequent pipetting steps. Optimization was carried out

for a volume of 18.7 l of feed solution (results not shown)

corresponding to the glucose mass flow of 13.3 g h1 l1.

Results of parallel cultivation are shown in Figures 4, 6

and 8. During batch growth, a maximum mean growth rate

() of 0.49+ 0.06 h1 was reached in the seven, millilitre-

scale, bioreactors. The initial glucose concentration of15 g l1 led to a final DCW of 7.2+ 0.5 g l

1 corres-

ponding to a biomass yield (YXS) of 0.48+ 0.03 g g1. The

DCW estimations based on measurements in the MTP

reader were characterized by a decreasing CV (Figure 5A).

The initially high CV of more than 10 % of the data from the

MTP reader is a result of the values corresponding to

the lower range of the preliminarily determined measure-

ment range. The CVs of DCW estimations derived from

DCUV650 measurements in the single-beam photometer are

comparable with the CVs of measurements in the MTP

reader at process times above 3 h (Figure 5B). Owing to

intermittent pH control with a frequency of 1 h1, pH

decreases to the value plotted in Figure 6 before the titration

agent is added. During batch growth, the reduction in pH/h

increases due to the increase in cell concentration. As a

result of acetate metabolism at the end of batch growth, pH

increases above the set point of 6.8 and reaches 7.05+ 0.05

at a process time of 4.8 h. During batch growth, a mean total

base volume (VBase) of 271+ 7 l was added to the cultures.

pO2 during batch growth was manually controlled at 20%

by increasing the impeller speed. After glucose depletion,

indicated by a sharp increase in pO2 at a process time of

3.76+ 0.04 h, the main metabolic by-product acetate was

consumed by the cells.

At the beginning of fed-batch operation, the glucose

mass flow was gradually increased in two steps to accom-modate the culture to intermittent feeding. At the end of

the process, a mean DCW of 36.9+ 0.9 g l1 was reached

in the seven bioreactors (Figure 4). During fed-batch growth,

the pH oscillated between 6.85 and 6.36 in all seven reactors

due to the intermittent pH control with a frequency of 3 h1

(Figure 6). In all samples taken during fed-batch growth, the

CV of pH did not exceed 1.1 % (Figure 7). The mean total

base volume added was 553.6+ 31.6 l.

From the beginning of fed-batch cultivation, pO2oscillates, owing to the intermittent feeding (Figure 8). After

the addition of glucose to the culture, pO2 decreases as a



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Figure 4 Parallel, automated, fed-batch cultivation ofE. coli K12 in seven,millilitre-scale, bioreactors in a mineral medium with glucose as the sole

carbon source

After exponential batch growth, a linear, stepwise increasing feeding profile

was started. Time course of DCW, substrate concentration (cS ; g l1) and

substrate mass flow (MS) are shown. Top to bottom: Reactors numbers 17.

Figure 5 Coefficient of variation of DCW measurements in the seven,

millilitre-scale, bioreactors

(A) CV of DCW atline measurements of the seven, millilitre-scale, bioreactors

in the MTP reader (DMTP650). (B) CV of DCW offline measurements of the

seven ml-bioreactors in the single-beam photometer (DCUV650).

result of the higher OUR (oxygen uptake rate; g l1 h1)

of the culture. After glucose depletion, pO2 increases to

its initial value. At the maximum glucose mass flow of

13.3 g h1 l1, a single glucose pulse results in a transitional

glucose concentration of 0.89 g l1. Although this value is

higher than the critical concentration of acetate formation

of 0.03 g l1 reported for E. coli [17], the culture is able

to metabolize the by-product during the feeding interval of

4 min, since no accumulation of acetate could be measured

until the end of the process.

After 14.6 h, the process started running into oxy-

gen limitation while the stirrer speed and volumetric oxygen

fraction in the inlet gas were kept constant. Assuming

(i) constant kLa by neglecting composition changes in theculture medium and (ii) negative influence of increase in

cell concentration, constant stirrer speed and volumetric

oxygen fraction in the gas inlet flow, guarantees a constant

OTR. Nevertheless, during intermittent feeding, the time

required to metabolize the added glucose decreases as

the cell concentration increases with ongoing process time

(Figure 9). The transitional decrease in pO2 concentration

after glucose addition is sharper and may thus lead to

transitional oxygen limitation at higher cell densities.

Monitoring pO2 enables the estimation of the OUR of

the culture. With an estimated kLa of 0.22 s1 of the impeller



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Figure 6 Parallel, automated, fed-batch cultivation ofE. coli K12 in seven,millilitre-scale bioreactors in a mineral medium with glucose as the sole

carbon source

Figure 7 CV of pH atline measurements of the seven, millilitre-scale,

bioreactors in the MTP reader with hydroplates

at 2600 rev./min [18] and a calculated oxygen solubility of

6.09 mg l1 for air saturation at 37C in the mineral medium

[19], the OUR profile depicted in Figure 10(A) can be

derived. During the constant feeding with a glucose mass

flow of 13.3 g l1 h1, the mean OUR gradually increases

from approx. 4 g l1 h1 to almost 6 g l1 h1. An increase

in oxygen uptake during constant feeding is a result of the

increase in maintenance of oxygen demand compared with

the oxygen demand due to growth. The lowest pO2 during a

feeding interval represents a maximum OUR of the culture.

Owing to the intermittent feeding profile, the maximum

OUR is more than 2-fold higher than the mean OUR of the

culture (Figure 10B).

Without taking the highest CV of the first three

DCW measurements in the MTP reader into account, the

overall mean CV is 4.4%. During 14.6 h, a mean DCW of36.9+ 0.9 g l

1 was reached. All relevant process data are

summarized in Table 2.

Since one of the major aims of bioprocess development

is the maximization of cell concentration per process time,

we compared the data reported here with the results of

Korz et al. [20], who developed a high-cell-density cultivation

process for E. coliby applying an exponential feeding profile.

Korz et al. reported an E. coli TG1 cell density of 35 g l1

biomass concentration after 22 h of exponential feeding

in a mineral medium. However, they started with a lower

inoculum density of approx. 0.1 g l1 resulting in a longer

batch phase of 12 h. Presuming an equal inoculum density ofapprox. 1.3 g l1 as in the present study, Korz et al. reached

their DCW of 35 g l1 after approx. 15 h, a time span

comparable with the process time in the present study. The

applied linear feeding strategy was thus able to compete

with the high cell density culture developed by Korz

et al [20].

Time course of pH as measured and controlled intermittently. For details on

titration frequencies, see the Process control subsection of the Materials and

methods section. Top to bottom: reactors numbers 17.



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Figure 8 Parallel, automated, fed-batch cultivation ofE. coli K12 in seven,millilitre-scale, bioreactors in a mineral medium with glucose as the sole

carbon source

Figure 9 Plot ofpO2

versus time during one feeding interval of 4 min for

three different process times and thus cell concentrations

Zero time corresponds to a feeding step at the feed mass flow

of 13.3 g l1 h1 . , tP (process time)=8.6 h (DCW=21.3 g l1); ,

tP =10.2 h (DCW=26.6 g l1); , tP =14.1 h (DCW=35.8 g l

1). For

clarity, broken lines have been added.

Table 2 Summary of mean process parameters of fed-batch cultivation in

seven, millilitre-scale, bioreactors

Process phase Parameter Mean value S.D. CV (%)

Batch DCW (g l1) 7.18 0.46 6.5

(h1

) 0.49 0.06 12.3YXS (g g

1) 0.48 0.03 6.5

pH 7.05 0.05 0.7

VBase (l) 271.2 7.0 2.6

pO2 increase 3.76 0.04 1.1

Fed-batch DCW (g l1) 36.87 0.91 2.5

VBase (l) 553.6 31.6 5.7

Time course of impeller speed, oxygen fraction in the inlet gas flow and pO2 .

Top to bottom: reactors numbers 17. Impeller [min1] means impeller speed

in rev./min.



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Figure 10 OUR in the seven ml-bioreactors

(A) Comparison of maximum OUR () with mean OUR (). Estimation

is based on a kLa of 0.22 s1 and an oxygen solubility of 6.09 mg l1 at air

saturation. (B) The ratio of maximum OUR to mean OUR during intermittent

feeding.

The higher maximum oxygen demand of a microbialculture during intermittent feeding could be satisfied by

the millilitre-scale bioreactors at an oxygen concentration

of 50% in the inlet gas flow. Our future work will focus

on the establishment of a closed-loop control of impeller

speed and oxygen concentration in the inlet gas flow, to

enable further prolongation of process times and thus

maximization of DCW. Moreover, the process control

software currently adapted is capable of controlling 48

bioreactors simultaneously.

Acknowledgments

We gratefully acknowledge the contribution of D. Ortlieb

(CFD Consultants, Rottenburg, Germany) in modelling the

hydrodynamics of the millilitre-scale bioreactor and the ex-

cellent technical assistance of M. Amann. This work

was partially supported by the Deutsche Bundesstiftung

Umwelt.

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Received 9 December 2004/15 April 2005; accepted 26 April 2005

Published as Immediate Publication 26 April 2005, doi:10.1042/BA20040197


miniature bioreactors for automated high-throughput bioprocess design

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