Biochemical Engineering Journal 39 (2008) 164176

Design and characterisation of a miniaalan

chemE 7JErdshieptem

Abstract

The establishment of a high productivity microbial fermentation process requires the experimental investigation of many interacting variables.In order to speed up this procedure a novel miniature stirred bioreactor system is described which enables parallel operation of 416 independentlycontrolled fermentations. Each miniature bioreactor is of standard geometry (100 mL maximum working volume) and is fitted with a magneticallydriven six-blade miniature turbine impeller (di = 20 mm, di/dT = 1/3) operating in the range 1002000 rpm. Aeration is achieved via a sintered spargerat flow ratestemperatureaddition, a nof individuaa function oB. subtilis Aspecific grow(1500 rpm, 1aeration rateachieved. Pr(1.5 L worki0.94 and 0.9for the paral 2007 Else

Keywords: M

1. Introdu

The deprocesses racting bioloengineeringrapid creati[1]. Theseand a smallple initial s

CorresponE-mail ad

1369-703X/$doi:10.1016/jin the range of 02 vvm. Continuous on-line monitoring of each bioreactor is possible using miniature pH, dissolved oxygen andprobes, while PC-based software enables independent bioreactor control and real-time visualisation of parameters monitored on-line. Inew optical density probe is described that enables on-line estimation of biomass growth kinetics without the need for repeated samplingl bioreactors. Initial characterisation of the bioreactor involved quantification of the volumetric oxygen mass transfer coefficient asf agitation and aeration rates. The maximum kLa value obtained was 0.11 s1. The reproducibility of E. coli TOP10 pQR239 andTCC6633 fermentations was shown in four parallel fermentations of each organism. For E. coli (1000 rpm, 1 vvm) the maximumth rate, max, was 0.68 0.01 h1 and the final biomass concentration obtained, Xfinal, was 3.8 0.05 g L1. Similarly for B. subtilisvmm) max was 0.45 0.01 h1 and Xfinal was 9.0 0.06 g L1. Biomass growth kinetics increased with increases in agitation and

s and the oxygen enrichment for control of DOT levels enabled max and Xfinal as high as 0.93 h1 and 8.1 g L1 respectively to beeliminary, scale-up studies with E. coli in the miniature bioreactor (100 mL working volume) and a laboratory scale 2 L bioreactorng volume) were performed at matched kLa values. Very similar growth kinetics were observed at both scales giving max values of7 h1, and Xfinal values of 5.3 and 5.5 g L1 respectively. The miniature bioreactor system described here thus provides a useful toollel evaluation and optimisation of microbial fermentation processes.vier B.V. All rights reserved.

iniature bioreactor; Parallel operation; Fermentation; On-line monitoring

ction

sign and optimisation of industrial fermentationequires the experimental investigation of many inter-gical and physical variables. Advances in metabolicand protein evolution techniques now enable the

on of large libraries of recombinant microorganismsare normally evaluated in parallel microwell culturesnumber of promising strains identified based on sim-

creens for product yield or activity [2]. For the chosen

ding author. Tel.: +44 20 7679 7942.dress: g.lye@ucl.ac.uk (G.J. Lye).

strains product synthesis can be further enhanced by study of cul-ture medium composition, nutrient feeding regimes and physicalvariables such as temperature, pH and dissolved oxygen levels.Here large numbers of parallel shake flask or stirred-bioreactorexperiments must be performed because of the number of exper-imental variables requiring investigation per strain. Finally, oncea production strain is identified, further experimental investiga-tion over a range of scales is necessary to establish the operatingboundaries and robustness of the process for validation purposes[2]. At this stage virtually all processes would be performed instirred bioreactors because of the dominance of this design inthe chemical and biopharmaceutical sectors.

Small scale bioreactor systems, that enable parallel and auto-mated operation of several fermentations simultaneously, have

see front matter 2007 Elsevier B.V. All rights reserved..bej.2007.09.001system for parallel microbiN.K. Gill a, M. Appleton b, F. Bag

a The Advanced Centre for Biochemical Engineering, Department of BioTorrington Place, London, WC1

b Bioxplore, 50 Moxon Street, Barnet, HertfoReceived 2 August 2007; accepted 3 Sture stirred bioreactorfermentationsz a, G.J. Lye a,ical Engineering, University College London,, UKre, EN5 5TS, UK

ber 2007

N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176 165

Nomenclature

CL

CpC*

dBdidTDOThi

kLamO2

OTRmaxOURmaxPttmVvvm

XfinalYX/O2

Greek lemaxpS

the potentiiments aretimes andlel miniatustirred tankThe key dedetail latersatisfy all thopment theinformation[2]. This isin microwebioreactortaken for hhowever, itmatched bing operatiprocess daunit operatautomatedautomatedsuch scalestion of theproduct.

In this work we report on the design, instrumentation andcharacterisation of a novel miniature stirred bioreactor system

umperaiore

six-bringachil opmassterisn ra, kLasub

hly rhowtionusin

sultsrodu

teria

hemconcentration of dissolved oxygen in fermenta-tion broth (kg m3)normalised oxygen concentrationsaturated dissolved oxygen concentration(kg m3)width of baffle (mm)diameter of impeller (mm)diameter of vessel (mm)dissolved oxygen tension (%)distance between centre line of impeller and baseof vesselvolumetric oxygen mass transfer coefficient (s1)oxygen required for cell maintenance(mol gDCW1 h1)maximum oxygen transfer rate (mmol L1 h1)maximum oxygen uptake rate (mmol L1 h1)

power input (W)time (s)mass transfer time, 1/kLa (s)volume (L)volumetric air flow per volume of broth per minutefinal biomass concentration (g L1)yield of biomass on oxygen (g mol1)

(maximallel oEach bsinglemonitoture isa nove

of biocharacaeratioficientand B.be higbeen sa funcresultsthat rebe repactor.

2. Ma

2.1. Cttersmaximum specific growth rate (h1)probe response time (s)superficial gas velocity (m s1)

al to increase the rate at which the necessary exper-performed thus reducing fermentation developmentcosts [3]. In recent years various designs of paral-re bioreactor systems have been reported includingbioreactors, bubble columns and shake flasks [47].sign features of many of these will be discussed in(Section 3.5). While each bioreactor design aims toe requirements for rapid fermentation process devel-

re is normally a trade-off between throughput and thecontent of the data obtained from each experimentmost clear when comparing related developments

ll fermentations [610] with the small scale stirreddesigns considered here. Irrespective of the approachigh throughput fermentation process developmentwill be important that advances in throughput are

y the creation of microscale downstream process-ons. In this regard the generation of quantitativeta from a number of microwell-based downstreamions has recently been reported [3,11,12] as has theoperation of linked process sequences [13,14]. Thelinkage of upstream and downstream operations atnow offers the potential for the integrated optimisa-

entire process from fermentation through to purified

The che(Dorset, Upurity availTOP10 pQnase (CHM[15], and Btion studiessolutions a

2.2. Minia

2.2.1. DesThe min

similar to cbioreactormaximumcontents ofhead plate,

Mixingcally drivedi/dT = 1/3)There wasRushton tucylindricaluniformly10 mm belThe wholenon-rotatinheadplate orotating meheadplate.working volume 100 mL) that can support the par-tion of 416 independently controlled bioreactors.actor is of standard geometry being agitated by alade miniature turbine impeller. Continuous on-lineand control of pH, dissolved oxygen and tempera-

eved using miniature steam sterilisable probes whiletical density (OD) probe allows on-line estimation

growth kinetics. Bioreactor oxygen mass transfertics have been studied as a function of agitation andtes giving a volumetric oxygen mass transfer coef-, of up to 0.11 s1. Parallel E. coli TOP10 pQR239tilis ATCC6633 fermentations have been shown toeproducible. In the case of E. coli TOP10 it has alson how biomass growth kinetics and yields vary asof agitation and aeration conditions. Finally initialg constant kLa as a basis for scale-up have shownobtained in the miniature bioreactor can accurately

ced in a conventional laboratory scale stirred biore-

ls and methods

icals and microorganisms

micals used in this work were obtained from BDHK) unless otherwise stated and were of the highestable. RO water was used for all experiments. E. coliR239, which expresses cyclohexanone monooxyge-O) under the control of an l-arabinose promotor. subtilis ATCC6633 [16] were used in the fermenta-. Cells were maintained as 40% (v/v) glycerol stock

t 80 C.

ture bioreactor design and instrumentation

ign of individual bioreactorsiature bioreactor was designed to be geometricallyonventional laboratory scale stirred fermenters. Eachconsisted of a borosilicate glass vessel (100 mLworking volume), allowing visual inspection of thethe bioreactor, and was sealed with a stainless steelas shown in Fig. 1.of the bioreactor was achieved by a magneti-

n, six-blade miniature turbine impeller (di = 20 mm,fabricated from PEEK (poly-ether ether ketone).

a distance of 20 mm between the centre line of therbine and the base of the vessel (di/hi = 1). Eightmagnets (3 mm in diameter and 2 mm long) were

distributed and embedded into a small PEEK discow the miniature turbine as shown in Fig. 1(A).

impeller structure was mounted onto a hollow,g stainless steel stirrer shaft that screwed into thef the bioreactor. This design eliminated the need forchanical seals and reduced the area occupied on theThe PEEK impeller assembly, which was magneti-

166 N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176

Fig. 1. Mechsingle miniatuof head plate.

cally drivenEach bioreremovablethe vesselthe miniatstainless stmore efficially regula0200 mL muse atmosping systemwas frequerate.

As showfour probesoxygen tendescribed iable on thea thermocothe thin wsample remwhich hadUK). Liqudles, securewas fitted wprevented land exhausScientific,

SchemB) ar2.2.2.

On-tinutedwitzuz ALtd.ticaanical drawing showing key design features and dimensions of are bioreactor: (A) cross section through bioreactor; (B) plan viewFurther details given in Section 2.2.1.

from below, had a stirring range of 1002000 rpm.

Fig. 2.probe; (Section

2.2.2.Con

facilitaAG, SBonad(HELfor opactor was also equipped with four equally spacedbaffles of width 6 mm (dB/dT = 1/10). Aeration ofwas via a narrow sparger located directly beneathure turbine. The sparger was fitted with a 15meel sinter to create finer gas bubbles and promoteent oxygen transfer. The air flow rate was manu-ted by a standard laboratory rotameter in the range

in1 (Fisher Scientific, UK) and it was possible toheric air alone or oxygen-enriched air via a gas blend-(HEL Ltd., UK). The gas flow rate to each reactorntly checked in order to ensure constant gas flow

n in Fig. 1(B), the headplate accommodated a total offor continuous on-line monitoring of pH, dissolved

sion (DOT), optical density (OD) and temperature asn Section 2.2.2. As a result of the limited space avail-

headplate a narrow opening in the centre alloweduple for temperature sensing to be located within

alled hollow stirrer shaft. All liquid additions andovals were via a single multi-port on the headplate

five openings closed by self-sealing septa (HEL Ltd.,id additions were made via sterile hypodermic nee-ly mounted by luer lock fittings. The exhaust gas portith a stainless steel, water cooled, condenser which

iquid loss from the medium by evaporation. The inlett gas lines were filtered through 0.2m filters (FisherUK).

sterilisableThe opt

and was siprobe, Fig.were each lone of the dother approsurement o(D2) was pted light apof this secoby the brotof emissiontor (D2) bythis work oreported. Awas generaminiature b

2.2.3. DesTo supp

was designtors and thLtd., UK),assembliesbase uniting a comatic diagram of in situ optical density probe: (A) dimensions ofrangement of light source and detectors. Further details given in

line instrumentation of individual bioreactorsous on-line monitoring of each bioreactor wasusing a miniature pH probe (Hamilton Bonaduzerland), polargraphic oxygen electrode (HamiltonG, Switzerland), a narrow K-type thermocouple

, UK) and a novel optical probe (HEL Ltd., UK)l density measurement. All probes were steam.ical probe, as shown in Fig. 2(A), had a PEEK bodyzed to fit each bioreactor headplate. The end of the2(B), had a white light source and two detectors thatocated in sealed glass tubes. The light source (LS) andetectors (D1) were positioned directly opposite eachximately 10 mm apart. D1 thus gave a direct mea-

f the optical density of the broth. A second detectorlaced at right angles to the direction of transmit-proximately 5 mm away from the beam. The outputnd detector measured the amount of light scatteredh. The light source had a sufficiently narrow angle

such that any light reaching the right-angle detec-direct transmission was considered negligible. In

nly the data from the optical density signal (D1) iscalibration curve of on-line and off-line OD values

ted for each type of fermentation completed in theioreactors.

ign and control of parallel bioreactor systemsort the individual bioreactors a modular base united which securely located four miniature bioreac-eir respective peristaltic pumps (RS Components

rotameters, electrical disc heaters and magnetic drive. Cables from the various probes plugged into thewhich housed all the associated electronics creat-pact unit with a footprint of 480 by 362 mm. Up

N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176 167

to four of these modular base units can be used in a singlesystem.

A custoallowed indture bioreaused for coreported toiments withbe conductwhich piecrequired tosteps with eThe end towhich is aperature, pparametersto be displaactor.

2.3. Chara

The oxytor was as[18]. Beforfresh de-iobrated betwand air, reconstant te(1000200atmospheriuid in themass transsured dissoresponse ti

Cp = 1tm

whereCp issured by thresponse tiperformedfor kLa det

2.4. Parall

Two difthe case osisted of 1NaCl (Sigm(SigmaAlsterilised acomponentantifoam (pature was mpositionedampicillin

bioreactor was inoculated with 2 mL (2%, v/v inoculum) of brothfrom a shake flask culture (100 mL in a 1 L shaken flask) grown

h on sha

lled a3PO4the c

m

ed aroup

ofluti

L1m (pn of2H2(40

6H21 Cwas

ingre wcom

ationculuflas

flask0 rpmat 6

3PO4both, agin 10vvm

xygeor 5

t sys(to sn coions)

naly

dditi.2.2)s ofof aere

aciare drriateicro

ibrates ofminith 1

weigm piece of PC-based software was written whichependent monitoring and control of up to 16 minia-

ctors. This was based on an architecture originallyntrol of automated chemical reactors that has beenbe easily programmed and simple to use [17]. Exper-

multiple operating conditions and set points coulded by generating an experimental plan that dictatede of equipment should be running and what it wasdo. Each experimental plan can consist of severalach step being carried out automatically in sequence.each step is specified by a terminating condition,

ny parameter defined by the user, e.g., time, tem-H, OD, DOT, etc. The software also enabled keysuch as pH, DOT, temperature and optical densityyed in real time on interactive graphs for each biore-

cterisation of bioreactor oxygen transfer rates

gen transfer capability of the miniature bioreac-sessed using the dynamic gassing out techniquee each experiment 10 g L1 NaCl was dissolved innised water and the dissolved oxygen probe cali-een 0% and 100% saturation by sparging nitrogen

spectively. All experiments were carried out at amperature of 37 C at predetermined stirrer speeds0 rpm) and aeration rates (12 vvm) using eitherc air or air enriched with oxygen. Assuming the liq-bioreactor was well mixed, the volumetric oxygenfer coefficient, kLa, was determined from the mea-lved oxygentime profiles accounting for the probeme according to Eq. (1) [19]:

p

[tm exp

(ttm

) m exp

(tp

)](1)

the normalised dissolved oxygen concentration mea-e probe at time t, tm equals 1/kLa and p is the probeme (18 s at 37 C). All gassing out experiments werein triplicate with the maximum coefficient of varianceermination being 6.1%.

el E. coli and B. subtilis fermentations

ferent microorganisms were used in this work. Inf E. coli TOP10 pQR239 the growth media con-0 g L1 each of tryptone, yeast extract, glycerol,aAldrich, Poole, UK) and 50 mg L1 ampicillin

drich, Poole, UK). Each miniature bioreactor wass a complete unit (121 C for 20 min) with all medias (apart from ampicillin) and 0.2 mL L1 of addedolypropylene glycol 2000). After cooling the temper-aintained at 37 C (0.2) via an electrical disc heaterbeneath the glass vessel. Filter sterilised (0.2m)was added immediately prior to inoculation. Each

for 14izontalcontro3 M H

Ingrowthpreparrate g100 mLof a so12.1 gantifoasolutioCaCl2metalsCoCl20.5 g Lactorcontainperatumediainoculv/v inoshakenshakeand 30trolled3 M H

Fortationsbetwee1 and 2with oof 30%four poditionsaeratiocondit

2.5. A

In ation 2sampleOD600tions wPharmhere aappropeach mthe calvolumfor 15once w

in pre-the same medium at 37 C and 200 rpm on a hor-ken platform (New Brunswick, USA). The pH wast 7 (0.1) by the metered addition of 3 M NaOH and.

ase of B. subtilis ATCC6633 chemically definededia was used. The media components werend sterilised (121 C for 20 min) in five sepa-s [16]. One litre of biomedia consisted of: (a)

a 200 g L1 solution of d-glucose, (b) 895 mLon of 11.2 g L1 (NH4)2SO4, 15.2 g L1 KH2PO4,

K2HPO4, 4.0 g L1 Na2HPO4, and 1.1 g L1olypropylene glycol 2000), (c) 2 mL of a 246 g L1MgSO47H2O, (d) 1 mL of a 147 g L1 solution ofO and (e) 2 mL of an acidified (pH 1) solution of traceg L1 FeSO47H2O, 5 g L1 MnSO4H2O, 2 g L1O, 1 g L1 ZnSO47H2O, 1 g L1 MoO4Na22H2O,uCl22H2O, 2 g L1 H3BO3). Each miniature biore-sterilised as a complete unit (121 C for 20 min)group (b) media components. After cooling the tem-as maintained at 32 C and the remaining groups ofponents were added aseptically immediately prior to. Each bioreactor was inoculated with 10 mL (10%,m) of an actively growing culture (100 mL in a 1 Lk) with an optical density of approximately 2. Theculture was grown on the same medium at 32 Con a horizontal shaken platform. The pH was con-

.8 ( 0.1) by the metered addition of 3 M NaOH and.

E. coli TOP10 and B. subtilis ATCC6633 fermen-tation rates in the miniature bioreactors were varied00 and 2000 rpm and aeration rates varied between(certain experiments with E. coli involved aeration

n-enriched air maintaining the DOT at a set point0%). Fermentations were performed in parallel on atem either under identical agitation and aeration con-how reproducibility) or under different agitation andnditions (to show parallel evaluation of fermentation.

tical techniques

on to the on-line optical density measurements (Sec-, cell growth was also monitored by taking brothup to 1 mL at regular intervals and measuring the

ppropriately diluted samples (1 in 2 to 1 in 20 dilu-used) off-line (Ultraspec 4000 spectrophotometer,Biotech, USA). All biomass concentrations reportedy cell weight (DCW) concentrations and were from

experimentally determined calibration curves fororganism and medium. The DCWs used to createion curve were determined in triplicate from knownfermentation broth. After centrifugation at 1300 rpm(Eppendorf AG, Germany), cell pellets were washed0 g L1 NaCl solution and dried at 100 C for 24 hhed and dried 2 mL Eppendorf tubes.

168 N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176

3. Results and discussion

3.1. Minia

The basbe as geombioreactorsture turbinquantitativand efficienpractice theto a minimwas complrates of upthe assembination, thedrive shaftstirrer. Thisthe impelleimpeller drto house su2000 rpm wdrive. Evenuniform diuid volume1000 rpm o

To facilcould be alaboratoryeach vesse8 bar befortest, in whshowed noabsence ofodically anfurther facunits eachand dissolvdensity proeach bioreathe smallesports for litinuous bioplace via asible to indbioreactors

3.2. Characapability

Since thmicroorganbioreactoruseful for cuseful critegassing ouand aeratio

nfluence of bioreactor operating conditions on oxygen uptake kineticsynamic gassing out experiments. From left to right: () 2000 rpm,- ) 2000 rpm, 1 vvm; (- - -) 1500 rpm, 1 vvm; () 1000 rpm,) 1000 rpm, 1.5 vvm; () 1000 rpm, 1 vvm. Experiments performedin 10 g L1 NaCl solution as described in Section 2.3.

n uptake is seen to increase with both increasing agitationratiower

tions1res

speeo hation

ioreactor oxygen mass transfer coefficient (kLa) as a function of stirrerd aeration rate with atmospheric air: () 1 vvm; () 1.5 vvm; () 2 vvm.f kLa calculated from the dynamic gassing out data shown in Fig. 3 usingSolid lines fitted by linear regression.ture bioreactor design and operation

is of the miniature bioreactor design was that it shouldetrically similar to conventional laboratory scaleas possible and be agitated by a standard minia-

e impeller. This was dictated by the desire to obtaine and scaleable data from each miniature bioreactort oxygen transfer during microbial fermentations. Inse constraints limited the volume of each bioreactor

um of 100 mL, to ensure that the miniature impelleretely submerged in liquid and maximum agitationto 2000 rpm could be achieved. In order to simplifyly of each unit, and to minimise any risk of contam-need for a rotating mechanical seal on the impellerwas overcome by opting for a magnetically drivenwas only possible by compromising on the design of

r and the incorporation of a flat disk at the end of theive shaft as shown in Fig. 1(A). This was necessaryfficient magnets to enable agitation at rates of up toithout decoupling the impeller from the magneticwith these modifications visual observation showed

stribution of gas bubbles throughout the entire liq-of each vessel provided that the agitation rate was

r above.itate parallel operation, four miniature bioreactorsssembled and sterilised simultaneously in a typicalautoclave. A hydrostatic pressure test indicted thatl could typically withstand an applied pressure ofe rupturing. An initial medium sterilisation and holdich DOT, pH and OD were monitored over 4 dayssigns of contamination. This was confirmed by theany colonies from medium samples withdrawn peri-d grown on nutrient agar plates at 37 C for 24 h. Toilitate parallel and unattended operation of multiplebioreactor was instrumented with pH, temperatureed oxygen probes as well as a novel on-line opticalbe (as described in Section 2.2.2). The small area ofctor head plate meant that it was necessary to sourcet available probes. In order to then have sufficient

quid additions, sampling and the possibility of con-reactor operation all liquid handling operations tookspecially designed multi-port. Finally, it was pos-

ependently monitor and control up to 16 miniatureusing custom-written, PC-based software.

cterisation of bioreactor oxygen transfer

e majority of industrial fermentations use aerobicisms the oxygen transfer capability of the miniatureis of great interest. Parameters such as kLa are alsoomparing different bioreactor designs and provide arion for scale-up. Fig. 3 shows a series of dynamict experiments performed at different stirrer speedsn rates for aeration with atmospheric air. The rate of

Fig. 3. Iduring d2 vvm; (2 vvm; (at 37 C

oxygeand aevaluesa funcas 0.11

Thestirrerseen tin aera

Fig. 4. Bspeed anValues oEq. (1).n rates. For each condition the corresponding kLae calculated using Eq. (1) and are plotted in Fig. 4 asof stirrer speed. The maximum kLa was determined.

ults in Fig. 4 show that kLa increases linearly withd over the range investigated. Stirrer speed is alsove a more significant effect on kLa than increasesparticularly at aeration rates above 1.5 vvm where

N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176 169

the improvement in oxygen transfer is minimal. Classically theimpeller power input and aeration rate have been correlated inthe literature using the well known vant Riet correlation:

kLa = K(

PgV

)

S (2)

where Pg/V is the impeller gassed power per unit volume, S isthe superficial gas velocity, K is a constant and and are expo-nents in the range of 0.4 < < 1 and 0 < < 0.7, respectively [18].For kLa measurements in ionic solutions, the reported values ofthe constants K, and were 2.0 103, 0.7 and 0.2, respec-tively. Although the vant Riet correlation was determined formuch larger vessels than used here, fitted with standard Rushtonturbine impellers, the stronger dependency of kLa on Pg/V (andhence stirrer speed) than on uS shown in Fig. 4 initially suggeststhat results for the miniature bioreactor are consistent with theestablished theory. At present the Power number of the miniaturebioreactor is not known and so absolute values of Pg/V cannotbe estimated.

3.3. Parallel fermentations and in-situ monitoring

3.3.1. Reproducibility of parallel E. coli fermentationsA key requirement of any parallel bioreactor system is that

cultivations in separate bioreactors should be highly repro-

ducible if performed under identical operating conditions. Fig. 5shows biomass growth kinetics and dissolved oxygen tension(DOT) profiles for typical parallel E. coli fermentations in thefour-pot miniature bioreactor system. These were carried out ata fixed impeller speed and aeration rate of 1000 rpm and 1 vvm,respectively which gave a relatively low kLa value of 0.04 s1(as measured in 10 g L1 NaCl). Considering the first fermen-tation shown in Fig. 5 (bioreactor B1), the entire fermentationlasted a total of 540 min with the end of the exponential cellgrowth phase occurring around 300 min. This coincided withthe point at which oxygen mass transfer limitation occurred andthe measured DOT reached zero. The particular strain of E. coliused here is known to have a high specific oxygen demand so thefact that the measured DOT reached zero at this kLa value is notsurprising [15]. The period of exponential growth was followedby an almost linear increase in biomass concentration for a fur-ther 160 min at which point the culture entered stationary phaseand ceased to grow. The biomass growth kinetics and dissolvedoxygen profiles for the other three bioreactors were very similarapart from the DOT profile in bioreactor B4 which reached zeroDOT somewhat earlier.

Table 1 shows the key kinetic parameters derived from theindividual fermentation profiles shown in Fig. 5. The maxi-mum specific growth rates were very similar, giving an averagevalue of 0.68 0.01 h1 as were the final biomass concentra-tions achieved where the average value was 3.8 0.05 g L1.

Fig. 5. Parall condperformed at e to foel batch fermentation kinetics of E. coli TOP10 pQR239 grown under identical1000 rpm and 1 vvm as described in Section 2.4. B1B4 refer to bioreactors onitions: () off-line biomass concentration; (-) DOT. Experimentsur, respectively.

170 N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176

Table 1Reproducibility of batch E. coli TOP10 pQR239 fermentation kinetics from four parallel fermentationsParameter B1 B2 B3 B4 BAV

max (h1) 0.68 0.66 0.68 0.69 0.68 0.01Xfinal (g L1) 3.7 3.8 3.8 3.8 3.8 0.05tDOT = 0 (min) 290 300 300 243 283 27.2OURmax (mmol L1 h1) 18.1 18.4 18.4 18.4 18.3 0.15Kinetic parameters derived from the fermentation profiles shown in Fig. 5. B1B4 refer to bioreactors one to four respectively, while BAV indicates the mean valuesof the kinetic parameters (error indicated represents one standard deviation).

The time at which the measured DOT reached zero was also com-parable in the majority of cases. Overall these results indicateexcellent reproducibility of the four-pot system.

At the point where the measured DOT first reaches zero, itcan be assumed that the maximum oxygen uptake rate (OURmax)and the maximum oxygen transfer rate (OTRmax) are equal. TheOURmax can thus be estimated from:

OURmax = XYX/O2

+ mO2X (3)

where is the specific growth rate at the time point when theDOT first reaches zero, X is the corresponding biomass concen-tration, YX/O2 is the yield of biomass on oxygen, which wastaken to be 1.92 g g1 [20] and m is the oxygen required for cellmaintenance which was taken to be 0.003 mol O2 g1 h1 [20].

Similarly, OTR can be estimated from:

OTR = kLa(C CL) (4)where C* is the saturated dissolved oxygen concentration, esti-mated to be 6.9 mg L1 [21] andCL is the actual concentration ofdissolved oxygen in the fermentation broth. At the point wherethe measured DOT first reaches zero, CL is also zero and so Eq.(4) reduces to:OTRmax = kLaC (5)

From Eq. (3) the average OURmax for the fermentationsshown in Fig. 5 was calculated to be 18.3 0.15 mmol L1 h1.In order to estimate the corresponding OTRmax value from Eq.(5), bioreactor kLa values were measured by further gassing outexperiments using spent biomedia containing the antifoam PPG

Fig. 6. Parallperformed atel batch fermentation kinetics of B. subtilis ATCC6633 grown under identical condi1500 rpm and 1 vvm as described in Section 2.4. B1B4 refer to bioreactors one to fotions: () off-line biomass concentration; (-) DOT. Experimentsur, respectively.

N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176 171

used during fermentation experiments. The presence of PPGis known to significantly reduce kLa values [22] and the mea-sured value of 0.03 s1 was not surprisingly 25% lower thanthe value measured in electrolyte solution. Using this lower kLavalue the calculated value of OTRmax was thus estimated to be18.2 mmol L1 h1 which is in excellent agreement with thecalculated OURmax values.

3.3.2. Reproducibility of parallel B. subtilis fermentationsIn addition to showing the reproducibility of parallel fermen-

tations of E. coli, a facultative anaerobe, the flexibility of theminiature bioreactor system was shown using parallel fermenta-tions of B. subtilis, a strict aerobe. Fig. 6 shows typical biomassgrowth kinetics and DOT profiles for four parallel B. subtilisfermentations in the miniature bioreactors. These fermentationswere carrie1500 rpm aof 0.06 s1(bioreactor660 min wiapproximareached zerphase andfermentatiominiature b

Table 2individual fspecific grovalue of 0.4tration achfor the DOTof variationto that deteexcellent re

In additvalues (as dlike B. subtunder steadrates must

dCLdt

= kLFor this

and YX/O2

Table 2Reproducibiliparallel ferme

Parameter

max (h1)Xfinal (g L1)tDOT = 0 (min)kLa (s1)Kinetic paramB1B4 refermean values odeviation).

respectively [16] and the contribution from cell maintenanceis negligible. Over a short time period dCL/dt is constant andthe kLa was calculated. This method gave a kLa value of0.062 0.001 s1 which is in excellent agreement with thatmeasured using the gassing out technique under the same oper-ating conditions (Table 2).

3.3.3. Correlation of on-line and off-line optical densitymeasurements

As described in Section 2.2.2, each miniature bioreactor wasequipped with a novel on-line OD probe (Fig. 2). This wasdesigned to facilitate continuous on-line monitoring of biomassgrowth kinetics which could be important given the small vol-ume of each bioreactor and the potential supervision of up to

n-linns. (

ation:of onations2.4.d out at a fixed impeller speed and aeration rate ofnd 1 vvm, respectively corresponding to a kLa value. Considering the first fermentation shown in Fig. 6B1) it can be seen that the entire fermentation lastedth the exponential phase of cell growth lasting untiltely 540 min. At this point the measured DOT againo at which point the culture rapidly entered stationaryceased to grow. As found with E. coli, very similarn profiles for B. subtilis were obtained in all fourioreactors.shows the key kinetic parameters derived from theermentation profiles shown in Fig. 6. The maximumwth rates were again very similar, giving an average5 0.01 h1 while the average final biomass concen-ieved was 9.0 0.06 g L1. The average time taken

level to reduce to zero was 565 33.7 min. The levelseen for the B. subtilis fermentations is comparable

rmined for the earlier E. coli work and again indicatesproducibility of the four-pot bioreactor system.

ion to using the gassing out method to determine kLaescribed in Section 3.2) in the case of a strict aerobeilis an oxygen mass balance can also be used, wherey state conditions oxygen uptake and consumption

balance [21], thus from Eqs. (3) and (4):

a(C CL) OUR (6)

particular medium and strain of B. subtilis CLhave been reported to be 6.8 mg L1 and 1.6 g g1

ty of batch B. subtilis ATCC6633 fermentation kinetics from fourntations

B1 B2 B3 B4 BAV

0.46 0.45 0.45 0.44 0.45 0.019.0 9.0 9.0 9.1 9.0 0.06

529 572 550 608 565 33.70.061 0.062 0.061 0.059 0.061 0.001

eters derived from the fermentation profiles shown in Fig. 6.to bioreactors one to four respectively, while BAV indicates thef the kinetic parameters (error indicated represents one standard

Fig. 7. Omentatiofermentity plotfermentSectione measurement of optical density in E. coli TOP10 pQR239 fer-A) Comparison of biomass growth kinetics from an individual() off-line optical density; () on-line optical density. (B) Par-

-line and off-line biomass concentration data from nine identical. Experiments performed at 1000 rpm and 1 vvm as described in

172 N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176

Fig. 8. Influe(-) DOT. Expcontrolled at 5

16 bioreacin Fig. 7(Afermentatiothat from ostages of thcalibratedues of theoff-line datthat the on-growth kinline biomaE. coli fermbetween ththe robustnand fermen

3.4. Inuecoli fermen

The initSection 3.3oxygen limoxygen uptof fermentnce of agitation and aeration rates and oxygen enrichment on batch fermentation kineterimental conditions: (A) 2000 rpm, 1 vvm; (B) 1000 rpm, 2 vvm; (C) 2000 rpm, 1 v0%. Fermentations performed as described in Section 2.4.

tor units in parallel by a single operator. As shown) the growth profile generated during a single E. colin by the on-line OD probe was virtually identical to

ff-line OD measurements, apart from during the finale culture. The OD reading of the probe was initially

using the sterilised culture medium. Calculated val-maximum specific growth rate from the on-line anda were 0.67 h1 and 0.68 h1, respectively indicatingline data can be used for quantitative analysis of celletics. Fig. 7(B) shows a parity plot of on-line and off-ss concentration data collected from nine identicalentations. There is seen to be excellent correlation

e two types of biomass concentration data indicatingess of the probe after repeated cycles of sterilisationtations.

nce of agitation and aeration conditions on E.tations

ial E. coli and B. subtilis fermentations reported inwere carried out at relatively low kLa values and soitations were observed in both cases. To explore theake requirements of the E. coli strain further, a seriesations were carried at higher agitation rates of 1500

and 2000 rhigher aeraconstant ator oxygen-the miniatu

Fig. 8 sprofiles undderived kinare summa

there is a s(0.750.94in all caselowest agittion 3.3.1growth ratin kLa valube 46.8 mmlimitationsreaching zvidual fermbioreactorsthat the DFor agitatigraphs Cics of E. coli TOP10 pQR239: () off-line biomass concentration;vm and DOT controlled at 30%; (D) 2000 rpm, 1 vvm and DOT

pm (aeration rate remained constant at 1 vvm) andtion rates of 1.5 and 2 vvm (agitation rate remained1000 rpm). Aeration was with either atmospheric airenriched air utilising the gas-blending capability ofre bioreactor system.hows examples of biomass concentration and DOTer the different agitation and aeration conditions. Theetic parameters from the entire series of experimentsrised in Table 3. For aeration with atmospheric airignificant increase in both the maximum growth rateh1) and final biomass concentration (5.15.6 g L1)s when compared to the values obtained with theation and aeration conditions used previously in Sec-(0.68 h1 and 3.8 g L1). These increases in celle are broadly in line with the measured increaseses. The largest value of OURmax was calculated tool L1 h1. In all cases, however, oxygen transferwere observed to remain with the measured DOT

ero at some point during the course of the indi-entations. In order to overcome this problem thecould be aerated with oxygen-enriched air such

OT was maintained at a constant set point value.on at 2000 rpm and aeration at 1 vvm in this case,and D in Fig. 8 show that the DOT could be

N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176 173

Table 3Variation of bioreactor oxygen mass transfer coefficient and E. coli TOP10 pQR239 fermentation kinetics as a function of agitation and aeration conditions includingthe use of oxygen enrichment to control DOT levels

Parameter

1000 rpm2 vmm (B)

2000 rpm 1 vmm (C) 2000 rpm 1 vmm (D)

Gas blending DOT = 30% DOT = 50%kLa (s1) 0.06 max (h1) 0.79 0.86 0.93Xfinal (g L1) 5.6 7.6 8.1OURmax (mm 27.5 ND ND(A)(D) corre cient and values derived from Fig. 3. N/D = not determined.

maintainedrespectivelmum growand the higWhile gasstages of fsary to imdensity cutors.

3.5. Comp

Severaldevelopmedesigns hatoward mecused in indmajority ooperation iments be pIn this resphighest degthe deck osmaller uniwhile the labe repeated

In termtors featurivalues. Stirworking vomany caseblending [2tain sufficiedate for anyactor blockare necessa

be achievedble of fed-breported kLminiature bpreviouslyavailable C12 stirred b

lumes of 35 mL. However, the oxygen transfer rates arecantly lower compared to other systems described in this

(0.0dditi

dese neeproclablelues6 vem wl wit0.0

f a sed bygooL) inon [nt kLomp

miy repigh tlogyherso fae thedbatcAgitation and aeration conditions

1000 rpm1 vmm

1500 rpm1 vmm

2000 rpm1 vmm (A)

1000 rpm1.5 vmm

0.04 0.06 0.08 0.060.68 0.83 0.94 0.753.8 5.1 5.3 5.6

ol L1 h1) 18.4 35.7 46.8 26.3spond to the fermentation profiles shown in Fig. 8. Oxygen mass transfer coeffi

and controlled at a minimum level of 30 or 50%y. In the absence of oxygen limitations the maxi-th rate of the culture increased further to 0.93 h1hest biomass concentration obtained was 8.1 g L1.blending can perhaps be avoided during the initialermentation process development it will be neces-plement it during the later stages when high cellltures must be attained in the miniature bioreac-

arison with other miniature bioreactors

miniature bioreactor systems for rapid bioprocessnt have been reported in recent years. A number ofve been investigated but the recent trend has beenhanically stirred bioreactors as these are most widelyustry for development and large scale operation. Thef the bioreactors have been designed with paralleln mind and there is a consensus that 10s of experi-erformed in parallel for the systems to be of value.ect the bioreactor block [23] currently gives theree of parallelisation, 48, but must be operated on

f a dedicated laboratory automation platform. Thets (110 mL) tend to be single use disposable itemsrger ones (100200 mL) like that described here canly steam sterilised.

s of agitation and aeration systems those bioreac-ng mechanical agitation tend to have the highest kLarer speeds necessarily increase dramatically as thelume of the bioreactors drop below about 10 mL. In

s non-standard gas-sparging impellers [23] and gas

ing vosignifisection

In aalso beand thferentis avaikLa vaup to 1a systeparallearoundbility oachievstrated(100 msumpticonstawere c

Thequicklallel, htechnoresearc

flask tincludthe Fe

Table 44,25, this work] have been implemented to main-nt oxygen transfer rates. The highest kLa reported toof the miniature bioreactors has been for the biore-

, achieving 0.4 s1 [23]. kLa values of this orderry to support the microbial cell densities that couldin approximately half of the designs that are capa-

atch or continuous operation. Betts et al. [25] havea values in the order of 0.1 s1 for a 10 mL stirredioreactor (the detailed design for this vessel has beendescribed by Lamping et al. [26]). The commerciallyellstation (Fluorometrix, USA) [27,28] allows up toioreactors to be operated in a single run with work-

Comparison oand conventio(based on aerParameter

Working voluStirrer speedAeration ratekLa (s1)Inoculum (%wmax (h1)Xfinal (g L1)Kinetic param1 s1).on to stirred vessels, small scale bubble columns haveigned in recent years, high-lighting the importanced for a variety of bioreactor systems to address dif-ess applications. Currently a 200 mL bubble columnfrom Infors, Switzerland and is reported to achieve

of up to 0.16 s1 [29,30], with the capacity to operatessels in parallel. Doig et al. [16] have characterisedith up to 48 miniature bubble columns operating inh working volumes of 2 mL, producing kLa values of6 s1. Although these systems lack the mixing capa-tirrer, sufficient aeration and some degree of mixing is

direct sparging. These researchers have also demon-d correlation with a laboratory scale bubble columnterms of oxygen transfer and volumetric power con-

31]. Cell cultivations were also scaled up based ona to bench scale stirrer bioreactor, and the resultsarable.niature bioreactor systems described herein arelacing the traditionally popular shake flask for par-hroughput operation, given the level of sophisticatedproviding large quantities of data. However, somehave attempted to enhance the design of the shake

cilitate some degree of online monitoring, theseRAMOS (HITECZang GmbH, Germany) [32] andh-pro (DasGip,Germany) [33] systems.f E. coli TOP10 pQR239 fermentations carried out in miniaturenal laboratory bioreactors using constant kLa as a basis for scale-upation with atmospheric air)

Miniature bioreactor Laboratory bioreactor

me (mL) 100 1500(rpm) 2000 1000(vvm) 1.00 0.67

0.08 0.08orking volume) 2.0 6.7

0.94 0.975.3 5.5

eters derived from Fig. 9.

174 N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176

In order to asses the reliability of miniature high throughputsystems as a tool for scale-down studies, it is critical to demon-strate their ability to mimic the conditions and productivity thatwould be expected at larger scales. A majority of the minia-ture systems that have been reported over the years have notdone this, however, some scale comparison studies have beencarried out comparing the growth and productivity of Bacil-lus subtillis RB50 and the productivity of riboflavin [34] in thebioreactor block [23] and a 7 L bioreactor. Betts et al. [25]have carried out more rigorous scale-up studies using a definedscale-up criterion based on constant power per unit volume giventhe geometric similarity between their vessel and a conventional

Fig. 9. Batchcarried out atscales: (A) ofmatched kLa vDoig et al. [15

bioreactor. Comparing the results of the miniature bioreactor andthe 7 L vessel concluded reasonable agreement in terms of max-imum specDNA.

Virtuallexpected foa number foptical denswitch fromprobes oncAll the desstrated featsupervision

3.6. Scale-bioreactors

While tcapable ofasses thesection theand thoselaboratorysel fitted wbe consideexperimentof 0.08 s1rates at thetation andscales Fig.biomass gring agreemat both scexpected ogrowth phference themaximumtrations actwo scalesvalues ovegives confiments in ththe convenment.

4. Conclu

arale) detoryted iniatu

n maandieveiore

ot bifermentation kinetics of E. coli TOP10 pQR239 fermentationsminiature (100 mL) and conventional laboratory bioreactor (1.5 L)f-line biomass concentration; (B) DOT. Experiments performed atalues as described in Section 2.4. Laboratory bioreactor data from].

A pvolumlaboraevaluathe mioxygetationbe achscale bfour-pific growth rates for E. coli DH5 producing plasmid

y all the designs feature the standard on-line probesr effective bioreactor monitoring and control though

eature novel probes for on-line monitoring of culturesity [23,27,28, this work]. The striking trend is the

standard probe technologies to fluorescent/opticale the bioreactor volume drops below about 10 mL.igns for which parallel operation has been demon-ure dedicated PC-based software necessary for theof multiple units.

up from miniature to laboratory scale

he miniature stirred bioreactor described here isautomated parallel operation, it is important to

scale-up potential of the results obtained. In thisrelationship between miniature bioreactor results

obtained using the same E. coli strain in a typicalscale 2 L bioreactor (1.5 L working volume ves-ith two top driven Rushton turbine impellers) will

red [15]. As a basis for predictive scale translations were initially performed at matched kLa valuesbased on direct measurements of oxygen transfertwo scales. Table 4 shows the corresponding agi-

aeration conditions. For E. coli cultures at the two9(A) shows that there was good agreement betweenowth kinetics while Fig. 9(B) shows the correspond-ent between measured DOT profiles. Experiments

ales were performed without gas-blending so thexygen limitations toward the end of the exponentialase are again seen. Given the 15-fold scale dif-re is excellent agreement between the calculatedspecific growth rates and final biomass concen-hieved (Table 4). Similar agreement between thewas determined for matched fermentations at kLar the range 0.060.11 s1 (data not shown). Thisdence that results obtained from parallel experi-e miniature bioreactors can be rapidly translated totional scales used for fermentation process develop-

sions

lel miniature bioreactor system (100 mL workingsigned to be geometrically similar to conventionalscale stirred bioreactors has been constructed andn this study. The oxygen transfer characteristics ofre bioreactor were first evaluated in terms of thess transfer coefficient, kLa, as a function of agi-aeration rates. Values as high as 0.11 s1 could

d comparable to those found in typical laboratoryactors. For identical fermentations performed in theoreactor system, excellent reproducibility between

N.K. Gill et al. / Biochemical Engineering Journal 39 (2008) 164176 175

parallel E. coli and B. subtilis fermentations was shown interms of the average calculated maximum specific growth rates(0.68 0.0 1 1tilis, respeccould be imby increasetation of gfrom miniaaccord wittor (1.5 L wat matchedscaleable dmakes thelel optimisexperimentneering enview to impactors.

Acknowled

The autructure Fu(SRIF) anto establisneering. FPhysical SLtd., in thdentship foare gratefuwork.

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thors would like to thank the UK Joint Infras-nd (JIF), the Science Research Investment Fundd the Gatsby Charitable Foundation for fundsh the UCL Centre for Micro Biochemical Engi-inancial support from the UK Engineering andciences Research Council (EPSRC) and HEL

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Design and characterisation of a miniature stirred bioreactor system for parallel microbial fermentationsIntroductionMaterials and methodsChemicals and microorganismsMiniature bioreactor design and instrumentationDesign of individual bioreactorsOn-line instrumentation of individual bioreactorsDesign and control of parallel bioreactor systems

Characterisation of bioreactor oxygen transfer ratesParallel E. coli and B. subtilis fermentationsAnalytical techniques

Results and discussionMiniature bioreactor design and operationCharacterisation of bioreactor oxygen transfer capabilityParallel fermentations and in-situ monitoringReproducibility of parallel E. coli fermentationsReproducibility of parallel B. subtilis fermentationsCorrelation of on-line and off-line optical density measurements

Influence of agitation and aeration conditions on E. coli fermentationsComparison with other miniature bioreactorsScale-up from miniature to laboratory scale bioreactors

ConclusionsAcknowledgementsReferences