Hybridoma growth and monoclonal antibody production in a dialysis perfusion system

Hybridoma growth and monoclonal antibody production in a dialysis perfusion system B. Amos, M. AI-Rubeai, and A. N. Emery B B S R C Centre f o r Biochemical Engineering, School o f Chemical Engineering, University o f Birmingham, Edgbaston, Birmingham, UK

Hybridoma cells were grown in perfusion culture using a stirred reactor within which a tubular membrane was suspended. Nutrient and product flows through the membrane to and from the culture environment occurred by diffusion processes alone. A mathematical model of the transfer and reaction process enabled both the characterization of a membrane mass transfer coefficient and the prediction of the maximum cell number achievable under set conditions. Steady states in cell concentration were observed for a range of perfusion rates and membrane areas. Steady states could be maintained for over 180 h without further addition of serum. Antibody was accumulated within the reactor to high concentrations, and at yields on both basal medium and serum that were many times those achieved in other forms of batch culture.

Keywords:Hybridomas;perfusion;dialysis;monoclonalantibody;metabolism;mathematicalmodel

Introduction In simple batch culture of animal cells, such as the hybridoma used for monoclonal antibody production, the rather low cell density normally achieved, typically 0.5-2 x 10 ~ cells ml-1, results in low concentrations of product and places a serious constraint on the use of this method for large-scale production. Continuous homogeneous perfusion systems have been recognized 1-6 for their advantages, such as increased yield on medium used, decreased labor costs through longer run times, easy monitoring and control of cell culture conditions, lack of nutrient and metabolite concentration gradients, and ease of scale-up. The use of membrane-based perfusion systems allows the retention of high-molecular-weight components (antibody product and serum proteins) as well as cells. Dialysis is a process whereby molecules up to a certain size are allowed to diffuse freely across a membrane in both directions. Larger molecules will not pass through the pores. Dialysis will not involve any net transport of mass unless there is a difference in mechanical or osmotic pressure across the membrane. Dialysis culture has been used extensively since 1896, and its applications have been reviewed by Schultz and Gerhardt.7 Address reprint requests to Dr. AI-Rubeai at the BBSRC Centre for BiochemicalEngineering,Schoolof ChemicalEngineering,Universityof Birmingham, Edgbaston,BirminghamB15 2TI', UK Received4 November1993;accepted 13 January 1994 688

Enzyme Microb. Technol., 1994, vol. 16, August

For large-scale production of monoclonal antibodies, such dialysis culture offers several potential advantages over other high-cell-density systems. Unlike most homogeneous systems and some of the more easily scalable heterogeneous systems (for example, macroporous microcarriers and gel bead entrapment), antibody is retained and is therefore recovered at very much higher concentrations, thus easing purification. There is no need to supplement the perfusate stream with (expensive) serum components because such high-molecular-weight species are retained, unlike the case in many perfusion systems. Dialysis culture can be conducted in ordinary reactors, and no special equipment, expensive consumables, or technically difficult operations are needed. Culture homogeneity allows simple global control of the reactor and assessment of the physiological state of the cells; moreover, easy access to the biomass for sampling and the relatively simple technology used may ease process validation. Many groups have grown hybridoma cells inside lengths of dialysis tubing suspended within medium, 8-11 and order of magnitude increases in cell concentration and antibody concentration have been observed. However, without providing good mixing in the cell compartment, the maximum usable volume with such a system is limited. One method of achieving good mixing is to turn the system inside out so as to have a large cell compartment mixed using conventional stirrers, and a small medium compartment, continuously fed from a large reservoir. Such systems have been used by Comer et al. 1 at up to the 1,000-1 scale. Densities in excess of © 1994Butterworth-Heinemann

Dialysis culture of hybridomas: B. Amos et al. I x 107 cells m l - 1 were achieved with viabilities greater than 90%, and increases in antibody concentrations over batch culture of between 2- and 15-fold for m o u s e - m o u s e hybridomas and 15- to 30-fold for h u m a n - h u m a n hybridomas were reported. H a g e r d o r n and Kargi 12 have used a similar system with a spiral-wound tubular m e m b r a n e giving substantial increases in cell density. O t h e r ways of using dialysis culture have recently been reported. Linardos et al. 13 have put a dialysis m e m b r a n e within a continuous reactor, increasing cell densities by up to 40%, and permitting the study of the effect of nutrient and toxic metabolite concentrations on steady-state growth and death rates. Kurosawa et aL 14 have placed a fixed bed of macroporous microcarriers within a dialysis membrane. The aim of this work is to provide a detailed description and justification of such an intensified cell culture system which allows the retention of both cells and products, and achieves e c o n o m y of medium utilization and reactor use. A mathematical model is described that can be used to correlate the physical factors affecting cell growth, and to predict performance u n d e r differing conditions.

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Materials and methods

Cell line and culture conditions TB/C3 is a mouse-mouse, NSl-derived hybridoma cell, a subclone of WC2 which produces antibody to the C2 region of human IgG. 15 Cells were maintained in RPMI 1640 medium with 5% newborn bovine serum (NBS). All culture processes took place using medium to which 0.5 g 1-1 Pluronic F68 and 2.5 g 1-1 meat peptone had been added. The latter can reduce the need for expensive serum and provides economic supplementation of basal medium components 16 when high cell densities are to be achieved.

Reactor description A 2-1 capacity, unbaffled, hemispherical-bottomed, top-driven, glass bioreactor of diameter 0.12 m (Setric Genie Industriel, Toulouse, France) was used. One 4-bladed 45° pitch impeller of width 0.06 m and blade height 0.07 m was mounted 0.08 m from the headplate. Headspace aeration at 100 ml min- 1 with agitation at 100 rev min - 1 was used to supply batch cultures without the need for sparging. For dialysis cultures, dissolved oxygen tension (DOT) was controlled at 50% air saturation by intermittent addition of oxygen, either sparged, or via 1 m of silicone tubing (diameter 5 × 10-3 m, wall thickness 1 x 10-3 m) suspended within the liquid. The pH was controlled at 7.0 by automatic addition of CO2 or 1 M NaOH.

Dialysis mode description Figure 1 is a schematic diagram of the reactor used for this study. Visking tubing, 0.037 m 2 (0.012 m diameter), 12,000 dalton molecular weight cut-off (Medicell International Ltd., London), was suspended in the fermenter in a coiled configuration supported by thick copper wire acting as a shapeable former. The wire was sealed in silicone tubing, and this was threaded through the dialysis tubing. Mediumwithout serum was circulated from a reservoir through the dialysis tubing. The reservoir was replenished constantly with fresh medium, the excess being forced by air pressure to a waste reservoir. A steady state was defined as three successive daily samples with less than 10% variation in cell density.

Dialyzing serum NBS was dialyzed against 50 volumes of phosphate buffered saline (PBS) overnight at 4°C; the PBS was then replaced and the dialysis continued for a further 24 h. A control aliquot was also kept at 4°C but not dialyzed. Glucose was measured at 0.71 g 1-1 in the nondialyzed serum and was not detectable in the dialyzed serum.

Assays Cell concentration was determined by hemocytometer count, viability being estimated by trypan blue exclusion. Glucose was assayed with a clinical glucose analyzer (Boehringer Mannheim, Lewes, UK). Lactate was measured enzymatically using lactate dehydrogenase coupled to NADH reduction (Sigma, Poole, UK). Ammonia was assayed using the method of Fawcett and Scott) 7 Antibody concentration was determined using a sandwich type ELISA.

Results and discussion A typical batch culture of TB/C3 in a p H and D O T controlled reactor (pH 7.0, D O T = 50%), achieved a maxim u m cell density of 0.78 x 106 cells m1-1 after 51 h, and there was no stationalTJ phase, the cell concentration decreasing to 0.22 × 10to cells m l - 1 in a further 72 h. The viability index 18 was 42.3 × 106 cells m l - 1 h - 1. This was calculated as:

t [xtl• +Xt2 I.

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- tl)

)

wherexa is the cell concentration at time ti. The final antibody concentration was 45 Ixg m l - 1, lactate concentration reached 29 mM, and final ammonia concentration was 1.5 mM.

Effect of membrane area in the dialysis reactor The effect of m e m b r a n e area on steady-state cell number was studied by setting up dialysis bioreactors with different tubing lengths. The standard length of dialysis tubing used was 1 m, giving a m e m b r a n e area of 0.037 m 2. For this

Enzyme Microb. Technol., 1994, vol. 16, August

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Papers experiment 0.75, 0.5, and 0.25 m of tubing, giving membrane areas of 0.027, 0.018, and 0.009 m 2, were used in consecutive runs. Steady states were obtained at a 0.5 v/v/d perfusion rate (see Figure 2); cell concentrations decreased linearly with tubing length (Figure 3) and were therefore apparently limited by transfer across the membrane under these conditions. The relationship between steady-state cell number and membrane area is not expected to be linear over the whole of its range: at zero membrane area the steady-state cell number should be zero (no nutrient transfer); however, the regression line does not return to the origin. At large membrane areas cell concentration should be limited by nutritional factors in the culture fluid. Even more m e m b r a n e area could be achieved using a longer piece of dialysis tubing, but this becomes impossible geometrically. During these steady states the proportion of dead cells was approximately 70%. A 3H-thymidine incorporation test showed that D N A synthesis was occurring during the steady states (results not shown). This suggests that the steady state must be a balance between cell division and cell death. As dead cells are retained within the reactor by the membrane, they accumulate until the rate of dead cell disintegration is equal to the rate of cell death. A1-Rubeai et al. 16demonstrated that cells remain viable in low serum concentrations because of continuing cell division, and not because of the extension of the survival of existing cells.

Effect o f perfusion rate on cell concentration

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In a membrane-limited system the steady-state cell concentration maintained is a complex function of medium perfusion rate. In order to study the effect of the rate of exchange of low-molecular-weight components on cell concentration, a range of steady states was achieved by feeding basal medium to the system at different fresh feed flow rates (see

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690

Enzyme Microb. Technol.,

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scribed above at a constant recirculation flow rate, using medium supplemented with peptone and pluronic. At a fresh feed flow rate of 0.5 1 d - 1, the steady-state cell concentration was 1.9 x 106 cells ml - 1. For a continuous culture without membrane retention at the same dilution rate, using the same medium and supplements, the steady-state cell concentration was 1.2 x 10°cells ml - 119 (it should be noted that serum was only added once in the reaction volume, at the start of the dialysis culture, in contrast to its continuous addition in continuous culture). This 30% improvement in cell yield on basal medium is a reflection of the energetic costs entailed in synthesizing biomass. Cell concentration increased with increasing perfusion rate, the increase tailing off at higher perfusion rates.

Dialysis culture of hybridomas: B. Amos et al. Table 1 Steady-state cell concentrations ( x 106 ml perfusion rates (d -1) Perfusion rate

1) at a range of

Cell concentration

0.5 1.0 1.5 2.0 2.5

1.9 2.8 3.4 3.9 4.2

Calculation o f membrane transfer coefficients To gain a better insight into the movement of nutrients through the system, and the physical factors affecting it, a mathematical description of the system was developed. Modeling of such systems was first attempted by Schultz and Gerhardt 7 and has also been performed by Hagerdorn and Kargi 12 and Szperalski et al. 20 whose differential equations give computer-based solutions for the unsteady states used in these studies. In the case of the steady-state conditions achieved in our system, however, simplified linear equations can be developed. Figure 5 shows the distinction between the recirculated volume Vc (of the circulation tube and recirculation loop) into which fresh substrate is fed and from which a product stream is bled and the reaction volume separated from it by the semipermeable membrane. A limiting substrate diffuses through the membrane and is consumed rapidly and irreversibly by the biomass in the reaction volume Vr. On the recirculating side the flow rate around the loop can be assumed to be high enough to assume an effectively constant concentration Sc of substrate in the mixing vessel and loop. The vessel is sufficiently well mixed that the fresh feed substrate (concentration Sf) can be considered to be instantly diluted, and, since the product stream is taken from this vessel, Sc is the concentration of substrate in that stream also. The substrate passes through the membrane due to the concentration gradient between the recirculated volume (Sc) and the reaction side (Sr). The membrane is characterized by an area Am and a transfer coefficient Kin. The latter strictly is an overall transfer coefficient for each side of the membrane, but, given the high flow velocities on each side, the film coefficients in the liquid streams can be ignored so long as the same process conditions are maintained. In the reaction volume Vr the substrate is consumed by the biomass at a specific rate Qs. In the steady state the biomass concentration is maintained at a fixed valuex. This is not to imply that there is no cell growth but rather that the net rate of (cell growth-cell death) is zero. Qs therefore represents consumption both for growth and for maintenance. Hence, at steady state, with no accumulation of substrate on either the recirculated side or the reaction side of the membrane, the consumption of substrate in the reaction side equals both the transfer through the membrane and the difference in substrate flux between the fresh feed and product streams (both at flow rate F), i.e.:

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F(Sf-Sc) (2) Am(Sc - Sr) The transfer coefficient for any substrate can therefore be calculated using only measurable parameters. Similar balances can of course be made for products, including possible toxic metabolites such as ammonia and lactate. Replacing the (S) terms by (P) terms and assuming such components are not found in the fresh feed stream, then one can readily show that the product transfer coefficients are given by: FPf gm (3) Am(Pc -Pr) The central assumption of such analysis is that the Qs and Qp terms are constant and independent of the Sr andPr values. The same assumption has been used elsewhere 21 to estimate cell numbers in heterogeneous reactors where direct measurements of cell numbers are difficult. Miller et al. 22,23 showed that Qs values for glucose and glutamine vary independently with their concentrations and with specific growth rate and pH for a hybridoma cell line in continuous culture. However, this present series of experiments was carried out at constant pH and DOT, and also with significant glucose concentrations in the product stream at all states, the fresh feed medium being supplemented with 2 g 1- 1 glucose. The validity of the assumption is further supported by the fact that the observed nutrient and cell concentrations measured over the full range of flow rates used match very well (see Table 3) those calculated using a specific uptake rate calculated for the lowest flow rate. The membrane transport coefficients measured for glucose and ammonia are identical with those given by Hagerdorn and Kargi 12 (see Table 2); the value for lactate is similar. Through a further simplified analysis it is possible to estimate the maximum cell number that a given reactor configuration can support. If it is assumed that the cell number in the reactor is limited by a substrate S that can pass freely through the membrane, then its concentration in

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E n z y m e M i c r o b . T e c h n o l . , 1994, v o l . 16, A u g u s t

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Papers Table 2 Comparison of membrane transfer coefficients from Hagedorn and Kargi 16 and those for this system. Data from Figure 4 were applied to equation (2). Units are (m h-1 × 10-3)

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Simulated values were based on K m values and Qs values calculated from the 0.5 I d-1 perfusion rate. Fr, Perfusion rate; X, cell concentration; Sr, glucose concentration in reactor; Sc, glucose concentration in circulation tank

the reaction volume Sr can be assumed to be rather low, certainly much less than Sc, so that, putting Sr = 0 into equation (2):

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A plot of ~/~against VFgives a straight line of slope K/xm with intercepts on the VFand 1Laxes of - VKand Kmrespectively. Applying this to the data in Table 1, a least-squares linear regression was used to calculate the equation of the line from which values forXm and Kwere calculated (see Figure 6): Xm = 5.88 x 106 cells m l - 1 and K = 1.05 h - 1. These equations have therefore given a prediction of steady-state cell concentration without knowing the identity of the limiting nutrient. Increasing the flow rate brings diminishing returns at values above that of K, so a maximum useful perfusion rate will be reached based on economic factors. Using the K value from the regression to calculate the Km for the limiting substance, a figure of 1.2 m h - ] x 10-3, was obtained, a value at the low end of the scale of those calculated and shown in Table 2. This suggests that the rate-limiting component is unlikely to be lactate or ammonia, which have larger transfer coefficients, though it may be glucose; the other likely possibility is glutamine. Glutamine is supplied by release from peptides in the peptone as well as in basal medium, and therefore its rate of use is not simple to assess. The maximum achievable cell concentration predicted for this system is, then, approximately 5.88 x 106 cells ml - 1, compared to the maximum we have achieved, which is 4 x 106 cells ml - 1 This model assumes that cell number is solely controlled by the level of a single low-molecular-weight component. However factors in the serum may affect the efficiency with which the cells utilize such a metabolite, and thus the cell concentration. To increase the cell concentration substantially from the present value would require large increases in perfusion rate. Increasing membrane area or using a

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Dialysis culture of hybridomas: B. Amos et al. different membrane type with a higher porosity, or molecular weight cutoff, will increase the rate of transfer of the limiting substance, and so also increase the steady-state cell concentration.

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Triplicate TB/C3 batch cultures of 10 ml volume in T-flasks were set up with RPMI plus peptone and Pluronic, with whole serum, and with dialyzed serum. Cell growth in dialyzed serum was higher than in nondialyzed serum (Figure 7). The viability index increased from 102.9 x 106 to 120.8 × 106 cell hours, and maximum cell density increased from 1.14 to 1.37 x 106 cells ml - ~. The increase in concentration noticed in dialyzed serum may be due to the removal by dialysis of a low-molecular-weight cytotoxic compound such as urea or uric acid. Low-molecular-weight components of the serum are seemingly not required in rich basal medium as much as high-molecular-weight components. Figure 8 shows that cells have been maintained in dialysis mode for 280 h without addition of serum, pH and D O T were maintained at 7.0 and 50% air saturation, respectively. The reactor was inoculated with 0.4 x 106 cells ml - l, and perfusion was started immediately at 0.51 d a y - ]. A steadystate cell concentration was achieved at 1.5 x 106 cells ml - 1 at 40 h. At 170 h the perfusion rate was increased to 1.0 1 day-~, and the cell concentration increased to 2.0 × 106 cells ml - 1. Subsequent periodic additions of 25-ml aliquots of NBS into the culture compartment led to temporary increases in steady-state cell concentration of up to 50%. After approximately 2 days the viable cell concentration in each case then returned to the original level. The reason for the increases was assumed to be due to components that were either of low molecular weight and so rap-

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idly diluted, or were short lived, or were rapidly sequestered. The bioreactor glucose concentration was zero during this period (Figure 8). In order to examine the possibility that the reactor referred to in Figure 8 was glucose-limited, another dialysis fermentation was then set up (Figure 9); the same medium was used as in the previous experiment, but with an additional 5.5 mmol 1- ~glucose in the feed, bringing the total for the feed to 16.6 mmol 1-1. The steady-state cell concentration at a feed rate of 11 day- 1was then 3.1 × 106 cells m l - 1, compared with 2.0 × 106 cells m l - 1 in the previous run (Fi~,ure8). The residual glucose concentration was 3 mmol 1- L At 192 h a 10% v/v solution of NBS in RPMI 1640 was added continuously at such a rate that 1% vN of serum was

Enzyme Microb. Technol., 1994, vol. 16, August

693

Papers added to the reaction volume daily. Volume was maintained at 1 1by removing the same volume daily when samples were taken. Under these conditions cell concentrations appeared to decrease. The effect then of adding additional serum appears to be less than the effect of removing cells. InFigure 8 it can be seen that, at a feed rate of 1 1d a y - 1, the bioreactor glucose concentration using the unsupplemented medium was below measurable limits, suggesting that the system was glucose-limited, although, as other nutrients were not measured, it may have been limited by others also. Twelve millimoles of lactate was present within the reactor, along with 0.8 mmol ammonia. These levels are half those occurring at the end of a batch culture of TB/C3. AI-Rubeai and Emery 24 have shown that addition of 15 mmol lactate in combination with 2 mmol ammonia to mid-exponential cells (typical levels for the end of batch culture of hybridomas according to Murdin et aL 25 reduces the rates of 3H-thymidine and 35S-methionine incorporation and MTT reduction in TB/C3 hybridoma cells by 21%, 25%, and 43% respectively. Hybridoma cells can also produce high-molecular-weight substances which are growth inhibitory. For example, Transforming Growth Factor [3 (TGF[3) is a 25-kD protein, produced by some hybridomas, which inhibits antibody production. Kidwel126 showed that antibody production was substantially reduced in hollow-fiber bioreactors with a 4-kD molecular weight cutoff, compared to using a 0.5-~xm pore size. TGF[3 was shown to be accumulated to inhibitory concentrations in a 4-kD molecular weight cutoff reactor, but not in that with a 0.5-1xm cutoff. Since the dialysis tubing used here has a nominal cutoff of 12 kD, such factors may be accumulated within our reactor.

Monoclonal antibody production In dialysis culture, antibody is retained within the system, and because the antibody molecule is relatively stable, it is accumulated to high concentrations. In Figure 8 the final antibody concentration is 1,440 mg m1-1, while in Figure 4 it is 1,200 mg ml - 1. These values are at least of the same order as those found in ascitic fluid (1-10 mg ml-1). 9 The specific antibody productivity range is 32-42 Ixg per 106 cells per day compared to 11-21 Ixg per 106 cells per day for a typical batch culture. Comparing the data with other data for this cell line in a variety of different culture devices and conditions, 19 the value is similar to that seen in a spin-filter device operated at high cell density, and higher than seen in batch, continuous, and the spin-filter system operated at low cell density. The reasons for high productivity in spin-filter systems would equally apply to this system: 1. Prolongation of the more productive G1 and S phases of the cell cycle; 2. Release of antibody from retained dead cells. Electron microscope preparations have shown that this cell line stores part of the antibody it synthesizes27 so the high fraction of dead cells seen in this culture will allow a lot of this antibody to be released.

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However, as the antibody is spending extended periods of time at 37°C, in a low-viability culture, antibody may be damaged by proteolytic action. For our dialysis system, then, the functions provided by high-molecular-weight components are either provided in excess, or released by the cells themselves (in which case cell density-dependent effects may be noticed), or may be provided by very stable macromolecules. Conclusions Dialysis culture is a simple economic method of increasing the volumetric productivity of standard reactor systems. Membrane transport performance can be characterized by a mass transfer coefficient that can be calculated from performance data. It is then possible to predict maximum cell numbers that can be maintained at a given operating condition. The use of dialysis culture allows substantial reduction in the use of serum, and this obviously reduces the medium costs of production considerably (see Table 4). The yield of antibody on basal medium is 1.4 times that in batch culture, and the yield of antibody on serum is 27 times that of a batch culture. The medium cost per gram of monoclonal antibody is therefore decreased by a factor of 5. Furthermore, the antibody is recovered as 25% of the total protein, whereas, in batch mode, the antibody is 2% of the total protein. Dialysis culture can produce order of magnitude increases in antibody concentrations over batch culture from similar increases in the number of viable cell hours. Its particular advantage of yielding a high concentration of product is especially beneficial when the cost of product purification is high, for example, where a product is normally present at very low concentrations. Although essentially a batch reaction is used, repeated draw and fill could be attempted and several harvests taken. More efficient membrane configurations, which permit increased maximal transfer rates, would have to be developed in order to allow scale-up while keeping the membrane in a manageable form. Hybridoma cell lines differ from each other in many ways, and will react differently to similar culture conditions. Experiments conducted with Chinese hamster ovary

Table 4 Comparison of a typical batch culture and dialysis cultures. Values are from Figure 4, except (*) from Figure 8. Medium costs are based on 1994 Gibco catalogue prices

Total medium used (I) Total serum used (I) Duration (h) M a x i m u m cell conc. ( × 106 cells ml 1) Viability index ( x 108 cells ml 1 h 1) Protein concentration (g I - 1)* A b concentration (g I 1) Yield Ab/basal medium (g I-1) Yield Ab/serum (g I - 1) Ab/total protein* £/g Ab

Batch

Dialysis

1.0 0.05 96 0.8 45.0 1.1 0.045

19.6 0.05 385 4.2 1080 5.0 1.2

0.045 0.90 2% 68.00

0.062 24 25% 15.07

Dialysis culture of hybridomas: B. Amos et al. (CHO) and insect cell lines have shown steady states of different cell densities from that seen with TB/C3 (results not shown).

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Acknowledgements

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T h e a u t h o r s gratefully a c k n o w l e d g e t h e s u p p o r t o f the S E R C B i o t e c h n o l o g y D i r e c t o r a t e . T h a n k s go to Miss S. C h a l d e r a n d Dr. Z. Z h a n g for their help.

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Nomenclature A membrane F K Pc Pf Pr Q Sc Sf Sr t V x

a r e a (m2) flow rate ( m 3 h - 1) t r a n s f e r coefficient ( m h - 1) product concentration (circulation volume) p r o d u c t c o n c e n t r a t i o n (feed) product concentration (reactor volume) specific rate ( m m o 1 1 0 - 6 v i a b l e cells h - 1 ) substrate concentration (circulation volume) s u b s t r a t e c o n c e n t r a t i o n (feed) substrate concentration (reactor volume) t i m e (h) v o l u m e ( m 3) cell c o n c e n t r a t i o n (10 6 cells ml - 1)

References 1

2

3

4 5

6 7 8 9

Comer, M., Kearns, M., Wahl, J., Munster, M., Lorenz, T., Szperalski, B., Behrendt, U. and Brunner, H. Industrialproduction ofmonoclonal antibodies and therapeutic proteins by dialysis fermentation. Cytotechnology 1990, 3, 295-299 Feder, J. Development of large scale mammalian cell perfusion culture system. In: Modem Approaches to Animal Cell Technology (Spier, R. and Griffiths, J., eds.). Butterworth-Heinemann, Oxford, 1987, pp. 454-482 Jan, D., Emery, A. and A1-Rubeai, M. Optimisation of spin-filter performance in the intensive culture of suspended ceils. In: Animal Cell Technology--Developments, Processes and Products (Spier, R., Griffiths, J. and Macdonald, J., eds.). Butterworth-Heinemann,Oxford, 1992, pp. 448-451 Pinton, H., Rabaud, J., Engasser, J. and Marc, A. Cytoflow--a new perfusion bioreactor for research and production. Biotech Forum Eur. 1991, 5, 344-347 Tokashiki, Y. and Arai, T. High density culture of mammalian hybridoma cells using a perfusion culture apparatus with multi-settling zones to separate cells from the culture medium. Cytotechnology 1989, 2, 5-8 Velez, D., Miller, L. and Macmillan, J. Use of tangential flow filtration in perfusion propagation of hybridoma cells for production of monoclonal antibodies. Biotechnol. Bioeng. 1989, 33, 938-940 Schultz, P. and Gerhardt, P. Dialysis culture of microorganisms. Bact. Rev. 1993, 33, 1-47 Adamson, S., Fitzpatrick, S., Behie, L., Gaucher, G. and Lesser, B. In-vitroproduction of high titre monoclonal antibody by hybridoma cells in dialysis culture. Biotechnol. Lett. 1983, 5, 573-578 Pannell, R. and Milstein, C. An oscillating bubble chamber for laboratory scale production of monoclonal antibodies as an alternative to ascitic tumors. J. Immunol. Methods 1992, 146, 43-48

13 14 15

16

17 18

19

20 21

22

23 24

25

26 27

Rijinders, A., van der Maaden, J., van Meel, F. and Olijve, W. An easy procedure for the production of serum-free monoclonal antibodies in dialysis tubing. In: Proceedings of the 4th European Congress on Biotechnology. Vol. 3 (Neijssel, O., van dcr Meer, R. and Luybrn, K. C. A., eds.). Elsevier, Amsterdam, 1987, p. 609 Sjorgren-Jansson, E. and Jeansson, S. Large scale production of monoclonal antibodies in dialysis tubing. J. Immunol. Methods 1985, 84, 359-364 Hagedorn, J. and Kargi, F. Coiled tube membrane bioreactor for cultivation of hybridoma cells producing monoclonal antibodies. Enzyme Micro& Technol. 1990, 12, 824-829 Linardos, T., Kalogerakis, N., Behie, L. and Lamontagne, L. Monoclonal antibody production in dialysed continuous suspension culture. Biotechnol. Bioeng. 1992, 39, 504-510 Kurosawa, H., Marki, H., Niebuhr-Redder, C. and Matsumura, M. Dialysis bioreactor with radial-flow fixed bed for animal cell culture. J. Ferm. Bioeng. 1991, 72, 41-45 Jefferis, R., Lowe, J., Ling, N., Porter, P. and Senior, S. Immunogenic and antigenic epitopes of immunoglobulins. 1. Cross reactivity of murine monoclonal antibodies to human IgG with immunoglobulins of certain species. Immunology 1982, 45, 71-77 Jan, D., Al-Rubeai, M. and Emery, A. Towards lower-cost medium formulation for intensive animal cell culture. In: Upstream and Downstream Processing in Biotechnology IIL Vol. 16 (Hyughebaert, A. and Vandamme, E., eds.). Royal Flemish Society of Engineers, Antwerp, 1991, pp. 1%26 Fawcett, J. and Scott, J. A rapid and precise method for the determination of urea. J. Clin. Path. 1960, 13, 156-159 Renard, J., Spagnoli, R., Mazier, C., Salles, M. and Mandine, E. Evidence that monoclonal antibody production kinetics is related to the integral of the viable cells curve in batch culture. Biotechnol. Lett. 1988, 10, 91-96 AI-Rubeai, M., Emery, A., Chalder, S. and Jan, D. Specific monoclonal antibody productivity and the cell cycle--comparisons of batch, continuous and perfusion cultures. Cytotechnology 1992, 9, 85-97 Szperalski, B., Geipal, F., Lorenz, T. U. B., Wahl, J. and Comer, M. Dialysis culture of hybridoma cells. Poster at the European Society for Animal Cell Technology llth Meeting at Wurzburg, May 1993 Lee, D., Gregory, D., Haddow, D., Piret, J. and Kilburn, D. High density BHK cultures using porous microcarriers. In: Animal Cell Technology: Developments, Processes and Products (Spier, R., Griffiths, J. and Macdonald, J., eds.). Butterworth-Heinemann, Oxford, 1992, pp. 480-487 Miller, W., Wilke, C. and Blanch, H. A kinetic analysis of hybridoma culture and metabolism in batch continuous suspension culture: Effect of nutrient concentration, dilution rate, and pH. Biotechnol. Bioeng. 1988, 32, 947-965 Miller, W., Wilke, C. and Blanch, H. Transient responses of hybridoma cells to nutrient additions in continuous culture. I: Glucose pulse and step changes. BiotechnoL Bioeng. 1989, 33, 47%486 A1-Rubeai, M. and Emery, A. Effects of nutrients and waste metabolites on growth, metabolism and protein synthesis in hybridoma. In: Production of Biologicals from Animal Cells in Culture (Spier, R., Griffiths, J. and Meigner, B., eds.). Butterworth-Heinemann, Oxford, 1990, pp. 70-72 Murdin, A., Thorpe, J., Groves, D. and Spier, R. Growth and metabolism of hybridomas immobilised in packed beds: Comparison with static and suspension culture. Enzyme Microb. Technol. 1989, 11, 341-346 Kidwell, W. Filtering out inhibition. Biotechnology 1989, 7, 462-463 Al-Rubeai, M., Mills, D. and Emery, A. Electron microscopy of hybridoma cells with special regard to monoclonalantibody production. Cytotechnology 1990, 4, 13-28

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