Characterisation of a novel miniaturised bubble column bioreactor for high throughput cell cultivation

Biochemical Engineering Journal 23 (2005) 97–105

Characterisation of a novel miniaturised bubble column bioreactor for high throughput cell cultivation Steven D. Doig, Anh Diep, Frank Baganz∗ The Advanced Centre for Biochemical Engineering, Department of Biochemical Engineering, University College London, Torrington Place, London WC1E 7JE, UK Received 23 June 2004; received in revised form 27 September 2004; accepted 26 October 2004

Abstract A miniature (2 ml) bioreactor has been developed for the high throughput cultivation of microorganisms. The microplate bubble column bioreactor (MBCB) consists of a static deep well microtiter plate with polyethylene frits (pore size 20 ␮m) inserted into the base of each of the 48 wells. Air was supplied at a controlled flow rate to the underside of each well and in this way the air passing through the pores of the frit formed bubbles in the liquid biomedium. A volumetric mass transfer coefficient for oxygen (kL a) of up to 220 h−1 could be achieved in this novel system. Since the system was not shaken, as is usually required with cell cultivations in microtiter plates, its performance was simpler to model and predict. A miniature optical oxygen probe was used to measure the dissolved oxygen concentration and it was found that the derived kL a values were proportional to the superficial gas velocity over the range 0–0.02 m s−1 . The usefulness of this new device was demonstrated by carrying out 25 parallel experiments to optimise the growth conditions of Bacillus subtilis ATCC6633, a strict aerobe, over a range of pH values and C/N ratios. © 2004 Elsevier B.V. All rights reserved. Keywords: Bioreactors; Bubble columns; Gas–liquid mass transfer; Fermentation; Mass transfer

1. Introduction During the development of a microbial cell cultivation process there are four key stages; strain selection, strain enhancement, process optimisation and scale-up [1]. Initially many wild type strains are screened for formation of the product of interest. These strains come from either established collections or can be isolated from the natural environment. The performance of the most promising strains identified by this initial screen is then enhanced using a variety of genetic and physiological techniques. Introduction of foreign genes into process friendly host strains (e.g. Escherichia coli), directed evolution, random mutagenesis, and optimisation of carbon flows can all enhance cell performance [2,3]. Quantitative process information, such as growth and product formation kinetics, is then determined to discover the optimal tempera∗

Corresponding author. Tel.: +44 20 7679 2968; fax: +44 20 7209 0703. E-mail addresses: [email protected] (S.D. Doig), [email protected] (F. Baganz). 1369-703X/$ – see front matter © 2004 Elsevier B.V. All rights reserved. doi:10.1016/j.bej.2004.10.014

ture, pH and growth media composition. Finally, the process is scaled-up from laboratory through to manufacturing scale. During this development process, many native and modified cell-lines are created and many operating conditions are considered and therefore large numbers (>100) of experiments are usually performed [3]. Since development time is key to commercial success, approaches that increase the rate at which these experiments can be carried out are of great value and therefore high throughput methodologies are of increasing interest [4]. The elements in an ideal high throughput approach are (a) experiments can be performed in parallel, (b) experiments can be operated at a small-scale, and (c) experiments can be automated [4]. The most commonly used cultivation vessel in process development is the shaken flask [5,6]. Erlenmeyer flasks (100–2000 ml), filled with low volumes of biomedium (10–25%) are shaken to promote mixing and mass transfer via surface aeration. Unfortunately shaken flasks cannot be easily automated and the number of simultaneous experiments is limited to several tens. Moreover, since they are difficult

98

S.D. Doig et al. / Biochemical Engineering Journal 23 (2005) 97–105

Nomenclature A C Cp Csat CPR kL a OUR t tm ug vvm V X YX/O2

area (m2 ) concentration (kg m−3 ) normalised concentration (kg m−3 ) saturation concentration (kg m−3 ) carbon dioxide production rate (mmol l−1 h−1 ) volumetric overall mass transfer coefficient (s−1 ) oxygen uptake rate (mmol l−1 h−1 ) time (s) 1/kL a (s) superficial gas velocity (m s−1 ) volumetric air flow rate (l l−1 min−1 ) volume (m3 ) biomass concentration (kg m−3 ) yield of biomass on oxygen (g g−1 )

Greek letters η density (kg m−3 ) µ viscosity (kg m−1 s−1 ) τp probe response time (s)

to instrument, the complexity of forthcoming data is limited [6]. As an alternative to the shaken flask, the use of shaken microtiter plates (microplates) is becoming more prevalent [8–10]. Operated in a similar way, these devices offer the simplicity of shaken flasks, but due to their standard footprint (82 mm × 125 mm) have the added advantage of their straightforward integration onto robotic platforms (e.g. http://www.tecan.com; http://www.perkinelmer.com) [4]. Since microplates are commercially available in a range of geometries ranging from 6 up to 1536 wells per plate (with respective filling volumes of 10 ml down to 10 ␮l), they also provide a high degree of parallelism whilst working with small volumes. Limitations with shaken microplates are the relatively low rates of oxygen mass transfer, complexities with instrumentation and scaling-up difficulties. Maximum oxygen transfer capacities of about 30 mmol l−1 h−1 (kL a of about 140 h−1 ) have been reported [11,12]. This relatively low kL a, when compared to 300–500 h−1 for a stirred tank bioreactor, can lead to oxygen mass transfer limitations and under such conditions growth and product formation can be adversely affected [6]. Situating the microplate on a shaking platform limits the level of instrumentation and monitoring that can be achieved. Although oxygen and pH probes have been developed for use in microplates [12,13], shaking needs to be stopped in order to take the readings and this complicates process automation. Finally, characterising the engineering environment in a well of a shaken microplate is complex. Mixing and mass transfer are determined by the geometry of the well, the liquid fill volume and by a vari-

ety of competing forces (interfacial, axial, gravitational and viscous) as well as by properties of the fluid itself (density and diffusivity) [6]. Therefore it is difficult to predict rates of oxygen mass transfer and mixing times and this can lead to complications when scaling-up the cultivation to stirred tank bioreactors. Rao and coworkers [14] have developed a miniaturised (2 ml working volume) stirred tank as an alternative to the shaken microplate. Since the vessel was static it was possible to instrument it with pH, DOT and biomass optical probes and thus the level of information generated was much greater. Although the system was able to support the growth of an E. coli culture, kL a values were only 44 h−1 at maximum and therefore oxygen limitation was a key problem. In this paper we propose the use of gas sparging, an alternative to shaking, as a means of promoting mixing and oxygen mass transfer for microbial cell cultivation in microplates. Microplates with porous membranes (frits) inserted into their base are commercially available and by applying a positive pressure to the underside of each well a flow of gas through the biomedium can be achieved. In this way oxygen can be delivered to a growing microbial culture whilst the engineering complexities associated with shaking are avoided. Although small-scale bubble column devices have been considered before (15 and http://www.inforsht.com), the volume of these vessels is 200 ml and therefore the current work represents a major reduction in scale (2 ml) and increase in experimental throughput (48 wells per plate). The objectives of this paper are to demonstrate the technical feasibility of this novel idea, termed the microplate bubble column bioreactor (MBCB), for the high throughput cultivation of Bacillus subtilis ATCC6633, with particular attention paid to the factors affecting oxygen mass transfer. B. subtilis is an aerobe and is commonly used in the biotechnology industry for the production of enzymes, antibiotics and foods [20]. In the presence of oxygen B. subtilis carries out aerobic respiration, however anaerobic respiration (reduction of nitrate or nitrite) is carried out when oxygen is limiting [21]. Fermentation cannot occur in a minimal medium with glucose as the sole carbon source [21].

2. Materials and methods 2.1. Chemicals and microorganism All chemicals used in this work were obtained from Sigma–Aldrich Chemical Co., Poole, Dorset, UK. B. subtilis ATCC6633 was obtained from the American Type Culture Collection (http://www.atcc.org). 2.2. Growth media A chemically defined growth medium was used throughout this work and was sterilised in five separate parts by autoclaving at 121 ◦ C for 20 min. One litre of me-

S.D. Doig et al. / Biochemical Engineering Journal 23 (2005) 97–105

dia consisted of, (1) 100 ml of a 200 g l−1 solution of dglucose, (2) 895 ml of a solution of 11.2 g l−1 (NH4 )2 SO4 , 15.2 g l−1 KH2 PO4 , 12.1 g l−1 K2 HPO4 , 4.0 g l−1 Na2 HPO4 , and 1.1 g l−1 antifoam (polypropylene glycol 2000), (3) 2 ml of a 246 g l−1 solution of MgSO4 ·7H2 O, (4) 1 ml of a 147 g l−1 solution of CaCl2 ·2H2 O, and (5) 2 ml of an acidified (pH1) solution of trace metals (40 g l−1 FeSO4 ·7H2 O, 5 g l−1 MnSO4 ·H2 O, 2 g l−1 CoCl2 ·6H2 O, 1 g l−1 ZnSO4 ·7H2 O, 1 g l−1 MoO4 Na2 ·2H2 O, 0.5 g l−1 CuCl2 ·2H2 O, 2 g l−1 H3 BO3 ). All solutions were made using deionised water. The concentration of antifoam was rather high, but was equally necessary in both stirred tank and bubble column vessels. Foaming during cultivation Bacillus sp. is well documented, however antifoam additions up to 1 g l−1 were not detrimental. 2.3. 5.5 litre scale fermentations A 7 l stirred tank bioreactor (Bioprocess Engineering Services, Charing, Kent, UK) was used with a fill volume of 5.5 l. The vessel had an aspect ratio of 3:1 (height/diameter), was fitted with four diametrically opposed baffles and was agitated by a top driven impeller with three six-bladed Rushton turbines (diameter = 6.27 cm). Air was supplied at a flow rate of 5.5 l min−1 (1 vvm). pH was monitored using a pH probe (Ingold Messtechnik, Urdorf, Switzerland), but was not controlled. Oxygen concentrations (DOT) were measured using a Clark type oxygen electrode (Ingold Messtechnik, Urdorf, Switzerland). The oxygen probe was calibrated using nitrogen gas (0%) and air (100%) and the probe response time was 15 s [15]. Exit gases were collected and channelled to a mass spectrometer (Prima 600, VG gas Analysis, Winsford, Cheshire, UK) and the oxygen uptake rate (OUR) and carbon dioxide production rate (CPR) were determined. RTDAS (Acquisition Systems, Guilford, Surrey, UK) software was used to collect and log pH, temperature, OUR and CPR. A 2 ml frozen stock (−80 ◦ C, 50% glycerol) of B. subtilis was used to inoculate 200 ml of biomedium in a 1000 ml conical shake flask. After 24 h incubation at 32 ◦ C with orbital shaking (300 rpm) this culture was then used to inoculate the stirred tank bioreactor. Samples were periodically taken and the biomass concentration was determined spectrophotometrically at 600 nm (U-2001, Hitachi, Tokyo, Japan). Dry cell weights (DCW) were calculated using a calibration curve of optical density against the weight increase created by filtering a known volume of fermentation broth through a pre-dried 0.2 ␮m filter paper. The error associated with this measurement was ±3% based on five independently repeated measurements made on a sample containing 3.5 g DCW l−1 . 2.4. Estimation of volumetric oxygen mass transfer coefficient Volumetric oxygen mass transfer coefficient, kL a, is defined by,

99

volumetric mass flux (g l−1 s−1 ) = kL a (s−1 ) × concentration driving force (g l−1 )

(1)

Oxygen mass transfer rates were determined by three methods: (a) the dynamic gassing out method [17], (b) by mass balance of the inlet and exit gasses, and (c) by mass balance from the growth rate under oxygen limitation [10]. Dynamic gassing out experiments were performed during cell cultivation between 350 and 450 min. The air supply was disconnected and due to the microbial activity the DOT dropped and once it had fallen to about 10% the air supply was reconnected. Taking into account the oxygen probe response time, the rate at which the DOT increased was used to calculate the kL a according to Eq. (2) [16].
    1 −t −t (2) tm exp Cp = − τm exp tm − τ p tm τp where Cp represents the normalised oxygen concentration measured by the probe, τ p the probe response time, tm equals 1/kL a, and t the time. An oxygen mass balance can be written as: dC = [kL a × (Csat − C)] − OUR (3) dt where C is the dissolved oxygen concentration, Csat the saturation concentration of oxygen in the biomedium (assumed to be 6.8 mg l−1 ). Over a short time period dC/dt is constant and kL a was calculated. The second mass balance was based on the rate of biomass growth (dX/dt) during oxygen limitation (assuming no cell death and negligible maintenance requirement), dX (4) = YX/O2 × kL a × (Csat − C) dt The yield of biomass on oxygen (YX/O2 ) was assumed to be constant under all conditions since B. subtilis could only grow via aerobic respiration. Variation associated with these measurements derived from either experimental errors (such as biomass measurements, poor mixing or evaporation of biomedium) or from systematic errors (such as significant maintenance energy). The former was estimated as ±12% based on five independently repeated experiments quoted at the 95% confidence interval. The latter was at most equivalent to 0.01 h−1 (a value taken for E. coli) and therefore could have introduced a maximum error of 2% [18]. 2.5. Microplate bubble column bioreactor (MBCB) Several filter-plates (microplates with porous membranes as bases to each well) were considered in an initial screen, but most were unsuitable since the commonly used thin paperlike membranes easily burst under pressure. Therefore, a microplate with porous frits was used since these were securely fitted and were not damaged by the upward flow of gas. Fig. 1 shows the set-up of the MBCB. The microplate was obtained from Innovative Microplate (Chicopee, MA, USA). Each of

100

S.D. Doig et al. / Biochemical Engineering Journal 23 (2005) 97–105

Fig. 1. Schematic layout of the MBCB. Humidified air is supplied via an air pump to the underside of each well of the MBCB. The air pressure was monitored and controlled with a manometer and a pinch valve. The working volume of the MBCB was 2 ml per well. A pierced SantopreneTM was used to minimise liquid losses and the flow rate of the exit gas was recorded using a bubble meter. The shaded wells are those used for detailed study in this work. Not drawn to scale.

the 48 wells was rectangular in cross-section and had a 2 mm thick polyethylene frit (20 ␮m pore size) fitted in the base. The working volume of each well was 2 ml and the whole system was operating within a heated and humidified cabinet (ISF-1-W, Adolf Kuhner AG, Birsfelden, Switzerland). Humidified air was pumped via silicone tubing to the base of each well and the air pressure was determined with a Utube manometer and was controlled with a pinch valve. The microplate was capped with a SantopreneTM mat (Innovative Microplate, Chicopee, MA, USA) and to allow gas flow through each well this was pierced with a fine bore needle and connected to a gas flow meter (custom made bubble flow meter). B. subtilis cultivations in the MBCB were inoculated with 10% (v/v) of an actively growing culture (100 ml in a 500 ml shaken flask) with an optical density of about 2–3 (dry cell weight, 0.8–1.3 g l−1 ). An amount of 100 ␮l samples were taken periodically for biomass (DCW) evaluation. The oxygen probe used in the 5.5 l bioreactor was not suitable for the MBCB due to its size and instead an optical probe with a 2 mm diameter was used. The oxygen probe

and data logging were described previously [19]. Briefly, the oxygen sensor (Ava-OXY, Knight Photonics, Leatherhead, Surrey, UK) consisted of a ruthenium complex immobilised in a sol–gel matrix onto the end of an optical fibre. Blue light (470 nm) from an LED was passed down the optical fibre and caused the ruthenium complex to fluoresce at 600 nm and this emitted light was measured by passing it back down a second optical fibre to a spectrometer. Oxygen quenched this fluorescence and thus this device could be used as an oxygen probe during dynamic gassing out experiments in the MBCB. The response time for this probe was 11 s.

3. Results and discussion 3.1. 5.5 litre scale cultivation of B. subtilis ATCC6633 Fig. 2 shows the biomass, DOT and exit gas profiles of a typical batch cultivation of B. subtilis. The final biomass concentration was between 11.5 and 12 g DCW l−1 and therefore the yield of biomass on glucose (YX/S ) was determined

S.D. Doig et al. / Biochemical Engineering Journal 23 (2005) 97–105

Fig. 2. Cell growth (
), DOT (—), oxygen uptake (- - -) and carbon dioxide production (—) rates of B. subtilis growing in 5.5 l stirred tank vessel using 1 vvm air and an impeller speed of 1100 rpm.

as 0.58–0.60 g g−1 . The cultivation lasted for 700 min and cell growth was exponential (maximum specific growth rate, µmax , was 0.44 h−1 ) for the first 550 min and then from 550 to about 600 min the rate of growth became linear due to oxygen mass transfer limitations (DOT = 0%). During the final 100 min the culture entered stationary phase and stopped growing. The oxygen limitation could not be avoided since the impeller was set to its maximum speed of 1100 rpm and increasing the aeration rate above 1 vvm resulted in uncontrollable foaming. During this period both the oxygen uptake rate (OUR) and carbon dioxide production rate (CPR) became approximately constant at 90 and 75 mmol l−1 h−1 , respectively. The yield of biomass on oxygen (YX/O2 ) was determined by mass balance (YX/O2 = (dX/dt)/OUR), as 1.6 g g−1 (taken as an average over the first 550 min of the fermentation). These results are similar to those of others. Celik and Calik [22] observed YX/O2 to be 1.4 g g−1 for B. subtilis NRS1125 growing on a glucose media supplemented with yeast extract and YX/S was 0.34 g g−1 . The lower yields are consistent with formation of a product, β-lactamase, in the cited work. Fig. 3 shows the effects of limiting the oxygen supply, by reducing the impeller speed, on growth rate and OUR. After 460 min the impeller speed was reduced to 500 rpm. At this point the biomass concentration was about 4 g DCW l−1 and the DOT immediately dropped to zero. During the following 130 min the impeller speed was held constant and the OUR remained approximately constant at 24 mmol l−1 h−1 and the rate of biomass increase was 1.42 g l−1 h−1 and therefore the estimated kL a was 113 h−1 based on the oxygen mass balance and 131 h−1 based on the oxygen limited growth rate. Between 590 and 690 min the impeller speed was increased to 800 rpm and similar measurements were made; the steady state OUR was 48 mmol l−1 h−1 and the linear growth rate was 2.28 g l−1 h−1 , giving kL a of 226 and 209 h−1 , respectively. Using these mass balance methods and the dynamic

101

Fig. 3. Cell growth, DOT (—) and oxygen uptake rate (- - -) of B. subtilis growing in 5.5 l stirred tank vessel under oxygen limitation: (
) biomass growth without oxygen limitation; () biomass growth under oxygen limitation.

gassing out method, the kL a was estimated during cell cultivation at several impeller speeds and the results are shown in Fig. 4. It was found that all three methods were within 15% of each other. 3.2. Characterisation of the microplate bubble column bioreactor Several factors were considered during the characterisation; bubble size, water losses due to evaporation and aerosol formation, oxygen mass transfer rates and gas hold-up. It seemed likely that these factors would be influenced by the air flow rate through the liquid biomedium, or the superficial gas velocity (ug ) defined as volumetric flow rate/crosssectional area. Fig. 5 shows the effect of increasing the air pressure ( p) applied to the underside of the porous frit on

Fig. 4. Comparison of different methods for the measurement of kL a in the 5.5 l stirred tank vessel. (×) Dynamic gassing out method, () mass balance of oxygen, (
) mass balance from linear growth rate. Insert, raw data generated from a dynamic gassing out experiments during cell cultivation at an impeller speed of 700 rpm.

102

S.D. Doig et al. / Biochemical Engineering Journal 23 (2005) 97–105

Fig. 5. Effect of applied pressure ( p) on superficial gas velocity (ug ) in the MBCB for a range of wells representing the diversity found across a 48 well plate: () well 1; (
) well 2; (䊉) well 3.

ug . The effect of p on ug was different for each of the three randomly chosen wells (shown in Fig. 1), but the variability shown in Fig. 5 represented the extremes found across all 48 wells. No loss of liquid occurred through the frit under any of the experimental conditions. For wells 1 and 2 the p at which air broke through the frit was about 8 mbar, whilst for well 3 air broke through at 15 mbar. Once breakthrough had occurred, the rate at which ug increased was linear, however the rate of increase was different in each well. Bubble sizes were estimated visually in these three wells and a wide range within each well was observed ranging from <1 mm up to 3 mm. The larger bubbles formed from air passing between the wall of the well and the porous frit and the smaller bubbles seemed to be formed from air passing through the pores in the frit itself. We concluded that variation in the construction of the wells, in particular the efficiency of the seal between the frit and the well wall, accounted for the well-to-well differences observed in Fig. 5 and in this wide bubble size range. Such variation is undesirable, since it would be ideal to supply air to all 48 wells of the MBCB at the same air pressure. However, during the initial characterisation work the wells were supplied with air individually and later in this paper we describe how the problem was minimised during parallel experimentation. In other work using small-scale bubble columns, the superficial gas velocity (ug ) were lower than those used in this research. The maximum ug used in this work was 0.02 m s−1 whilst using a 200 ml vessel the highest ug was 0.006 m s−1 [15]. We used higher ug since we suspected most of the dispersed gas would escape from the vessel without oxygenating the liquid medium. Due to the large volumes of air that were passed through the liquid biomedium (the volumetric flow rate ranged from 0 to 75 vvm) excessive water loss, via evaporation or aerosol formation, was a major concern given the small-scale of operation. To minimise evaporative losses the incoming air was humidified to >85% saturation. Aerosol formation was reduced in two ways. First by fitting a SantopreneTM cap pierced by a narrow bore needle to al-

Fig. 6. Growth of B. subtilis in the MBCB compared to that in the 5.5 l stirred tank vessel: (
) MBCB, at a ug of 0.02 m s−1 ; () 5.5 l stirred tank, data taken from Fig. 2.

low gases to exit and second by adding antifoam PPG2000 to the biomedium. In this way the volumetric losses from the MBCB (aerated for 5 h at a ug of 0.015 m s−1 ) were reduced from 41 to 69% (v/v) down to between 6 and 17% (v/v). Such water losses were no worse than observed with shaken microplates [8]. Fig. 6 shows a typical growth curve of B. subtilis in the MBCB and compares it to the 5.5 l stirred tank cultivation. Well 3 was used in this experiment with a p of 30 mbar so that ug was 0.01 m s−1 . The final biomass concentration was 10.4 g DCW l−1 and between 570 and 735 min the oxygen limited linear biomass accumulation rate was 1.72 g l−1 h−1 and therefore the estimated kL a was 158 h−1 . Using the three wells from Fig. 5, cell cultivations were performed over a wide range of ug and from the oxygen limited growth phases the kL a values were calculated and are plotted in Fig. 7. Although there was significant variation, the general trend of

Fig. 7. Effect of superficial gas velocity (ug ) on the oxygen volumetric mass transfer coefficient (kL a) and gas hold-up in the MBCB. (
) kL a from linear growth rate, (䊉) kL a from dynamic gassing out, (×) gas hold-up. The solid line shows the fitted correlation to linear growth rate data, kL a = 0.65ug 0.6 . Data in this figure represents values obtained from all three wells and the values are not averaged.

S.D. Doig et al. / Biochemical Engineering Journal 23 (2005) 97–105

increasing kL a with increasing ug is clear. Also shown in Fig. 7 are kL a values determined by the dynamic gassing out method using the fluorescent oxygen sensitive optrode. These values were about 20–30% lower than those calculated using the oxygen limited growth rate, but the trends are very similar. It is likely that the difference between the two values was due to inefficient mixing in the liquid phase and therefore positioning of the optrode was vital. Unfortunately, the position of the optrode was limited to zones in the wells away from bubble streams since when air bubbles settled on the optrode tip spurious kL a values were measured. Therefore the oxygen limited growth rates associated kL a were taken as being more accurate. Many simple correlations exist for bubble columns in which kL a is found to be proportional to the superficial gas velocity, for example, kL a = 0.32(ug )0.7 [23]. Using the linear growth rate derived data, a plot of ln(kL a) versus ln(ug ) (not shown) was used to determine the exponent of 0.6 and the correlation, kL a = 0.65(ug )0.6 , was fitted by minimising the regression coefficient. This correlation is also shown in Fig. 7. Using this correlation the problem of well-to-well variability in ug at a given p, mentioned above, was further quantified. The key reason that variability in ug is important is through its effect on kL a. At a p of 35 mbar, the range of ug was between 0.015 and 0.02 m s−1 and therefore the concomitant range of kL a was 188–223 h−1 (20% variation). We believe that this level of variation is manageable considering the small-scale of operation. Gas hold-up [volume of gas/(volume of gas + liquid)] was dependent on the flow rate of gas through each well and the nature of the liquid. The gas hold-up in an air/biomedium (containing 1 g l−1 PPG 2000) system, measured from the increase in the liquid height when gassed, was estimated between 2 and 11% as shown in Fig. 7. The gas hold-up increased with increasing ug , and although the measurements were highly variable, it suggests that the two were closely related. In conventional scale bubble columns the gas holdup, bubble size and mass transfer rates are inter-related. The coalescent properties of the liquid, the superficial gas velocity as well as the sparger design determine the bubble size distribution and gas hold-up [23]. Since antifoam containing media, such as that used in this work, are highly coalescent, it might be expected that bubble coalescence would be a critical factor governing oxygen transfer rates. However, since the MBCB is small and the height of the liquid was approximately the same as the vessel diameter, these complications are reported to be negligible and thus were not considered in this work [24].

103

Fig. 8. Use of the MBCB in a parallel cultivation experiment: influence of pH and C/N (mol mol−1 ) ratio on µmax .

10% (v/v) of a standard B. subtilis culture grown overnight. The initial pH and media compositions were varied across the wells, but the temperature was kept constant. The initial pH was set at either 6.0, 6.5, 7.0, 7.5 or 8.0 and the glucose to (NH4 )2 SO4 ratio was varied to give C/N ratios of 0.94, 1.89, 3.77, 5.66 and 7.54 (mol C mol N−1 ) (where the (NH4 )2 SO4 concentration was kept constant at 5 g l−1 ). This type of optimisation is common during fermentation development, and allows for rapid identification of the optimal operating conditions. Two sets of data were captured during this experiment; specific growth rate (µ) and the yield of biomass on carbon source (YX/S ) and the results are shown in Figs. 8 and 9. It is clear that extremes of pH (6.0 and 8.0) the growth rate decreased, but it can be seen that between 6.5 and 7.5 there was a broad optimum. Interestingly the C/N ratio also affected µ where the general trend was that higher concentrations of glucose lowered the growth rate. The highest growth rate observed in the MBCB was 0.42 h−1 , which is only slightly lower than the value of 0.44 h−1 ob-

3.3. Screening cell performance in the MBCB using a simple experimental design In order to demonstrate the utility of the MBCB in high throughput assessment of microbial performance, an experiment was carried out in which 25 individual cultivations were performed in parallel. Each well was inoculated with

Fig. 9. Use of the MBCB in a parallel cultivation experiment: influence of pH and C/N (mol mol−1 ) on YX/S .

104

S.D. Doig et al. / Biochemical Engineering Journal 23 (2005) 97–105

served in the stirred tank bioreactor. As the C/N ratio was increased from 0.94 to 3.77 mol mol−1 , YX/S remained approximately constant at 0.52–0.6 g g−1 . Little differences in yield were observed at the various pH values. However, above this critical ratio the yield decreased rapidly were presumably glucose was in excess and nitrogen became the limiting nutrient. Using this experimental design the optimum conditions of pH 6.5–7.5 and a C/N ratio of 1.89 were determined.

4. Conclusions In this work we have provided a novel alternative to shaken vessels for small-scale and high throughput cell cultivation—the microplate bubble column bioreactor. This device, operating at 2 ml scale in microtiter plate format, was successfully used for the parallel cultivation of the aerobic bacterium B. subtilis and final biomass concentrations of up to 10.5 g DCW l−1 were achieved. Because oxygen transfer occurred via gas bubbling through the liquid biomedium, rather than surface aeration, volumetric mass transfer coefficients for oxygen of up to 220 h−1 were calculated and this is an improvement over the more conventional shaken microplate and other miniaturised vessels [14]. However, instrumentation of the MBCB with dissolved oxygen and pH probes, as achieved in the work of Kostov et al. [14] has not yet been achieved. We expect that similar instrumentation of an individual well in the MBCB should be no more difficult, but that instrumentation of each well of an entire microplate will be a significant technological challenge. We believe that many of the problems encountered during cell cultivation at microplate scale (200 ␮l to 5 ml) are due a lack of understanding of the engineering environment and the effect that it has on the microorganism. In particular, the oxygen mass transfer mechanism is very complex in shaken small-scale vessels and the effect that oxygen limitation can have during cell cultivation is undesirable for growth and product formation. As the MBCB was aerated and agitated by rising bubbles, rather than shaking, mass transfer processes are simpler to understand and rates easy predict. It was also possible to instrument the device with an oxygen probe and therefore a better definition of oxygen mass transfer was gained. We expect this new approach to small-scale, high throughput cell cultivation will provide better quality and more reliable information during screening of cellular performance particularly for high cell densities exhibiting high oxygen uptakes rates.

Acknowledgements The authors would like to thank the UK Joint Infrastructure Fund (JIF), the Science Research Investment Fund (SRIF) and the Gatsby Charitable Foundation for funds to establish

the UCL Centre for Micro Biochemical Engineering. Financial support for the research work presented in this paper is acknowledged from the UK Engineering and Physical Sciences Research Council (EPSRC).

References [1] S.D. Doig, F. Baganz, G.J. Lye, High throughput screening and process optimisation, in: C. Ratledge, B. Kristiansen (Eds.), Basic Biotechnology, 3rd ed., Cambridge University Press, 2004. [2] J.D. Keasling, Gene-expression tools for the metabolic engineering of bacteria, Trends Biotechnol. 17 (1999) 452–460. [3] M. Chartrain, P.M. Salmon, D.K. Robinson, B.C. Buckland, Metabolic engineering and directed evolution for the production of pharmaceuticals, Curr. Opin. Biotechnol. 11 (2000) 209– 214. [4] G.J. Lye, P. Ayazi-Shamlou, F. Baganz, P.A. Dalby, J.M. Woodley, Accelerated design of bioconversion processes using automated microscale processing techniques, TIBTECH 21 (2003) 29– 73. [5] U. Maier, J. Buchs, Characterisation of the gas–liquid mass transfer in shaking bioreactors, Biochem. Eng. J. 7 (2001) 99–106. [6] J. Buchs, Introduction to advantages and problems of shaken cultures, Biochem. Eng. J. 7 (2001) 91–98. [8] S.D. Doig, S.C.R. Pickering, G.J. Lye, J.M. Woodley, The use of microscale processing technologies for the quantification of biocatalytic Baeyer–Villiger oxidation kinetics, Biotechnol. Bioeng. 80 (2002) 42–49. [9] W. Minas, J.E. Bailey, W. Duetz, Streptomyces in micro-cultures; growth, production of secondary metabolites and storage and retrieval in the 96-well format, Antonie van Leeuwenhoek 79 (2000) 297–305. [10] W.A. Duetz, L. Ruedi, R. Hermann, K. O’Connor, J. Buchs, B. Witholt, Methods for intense aeration, growth, storage and replication of bacterial strains in microtiter plates, Appl. Environ. Microbiol. 66 (2000) 2641–2646. [11] R. Hermann, M. Lehmann, J. Buchs, Characterisation of gas–liquid mass transfer phenomena in microtiter plates, Biotechnol. Bioeng. 81 (2002) 178–186. [12] G.T. John, I. Klimant, C. Wittmann, E. Heinzle, Integrated optical sensing of dissolved oxygen in microtiter plates: a novel tool for microbial cultivation, Biotechnol. Bioeng. 81 (2003) 829– 836. [13] S. Weiss, G.T. John, I. Klimant, E. Heinzle, Modeling of mixing in 96-well microplates observed with fluorescence indicators, Biotechnol. Prog. 18 (2002) 821–830. [14] Y. Kostov, P. Harms, L. Randers-Eichhorn, G. Rao, Low cost microbioreactor for high-throughput bioprocessing, Biotechnol. Bioeng. 72 (2001) 346–352. [15] D. Weuster-Botz, J. Altenbach-Rehm, A. Hawrylenko, Process engineering characterisation of small-scale bubble column for microbial process development, Bioprocess Biosyst. Eng. 24 (2002) 3–11. [16] I.J. Dunn, A.J. Einsele, Oxygen transfer coefficients by the dynamic method, J. Appl. Chem. Biotechnol. 25 (1975) 707–720. [17] P.M. Doran, Bioprocess Engineering Principles, Academic Press, London, 1998. [18] S.J. Pirt, Principles of Microbe and Cell Cultivation, Blackwell Scientific Publications, London, UK, 1975. [19] S.R. Lamping, H. Zhang, B. Allen, P. Ayazi Shamlou, Design of a prototype miniature bioreactor for high throughput automated bioprocessing, Chem. Eng. Sci. 58 (2003) 747–758. [20] C.R. Harwood, Bacillus, Plenum Press, New York, 1989.

S.D. Doig et al. / Biochemical Engineering Journal 23 (2005) 97–105 [21] M.M. Nakano, P. Zuber, Anaerobic growth of a strict aerobe (Bacillus subtilis), Annu. Rev. Microbiol. 52 (1998) 165–190. [22] E. Celik, P. Calik, Bioprocess parameters and oxygen transfer characteristics in ␤-lactamase production by Bacillus species, Biotechnol. Prog. 20 (2004) 491–499.

105

[23] J.E. Bailey, D.F. Ollis, Biochemical Engineering Fundamentals, McGraw-Hill International, Singapore, 1986. [24] H.W. Blanch, D.S. Clark, Biochemical Engineering, Marcel Dekker, New York, 1997.