Bioreactors or fermenters for the cultivation of cells or microorganisms in suspension on an industrial scale

Abstract

To provide a bioreactor or fermenter for the culturing of cells or microorganisms in suspension in a liquid medium on an industrial scale.SOLUTION: The present invention relates to a bioreactor or fermenter comprising: a vessel containing the culture in a liquid medium having a determined filling height; a stirrer provided in the vessel to stir the liquid medium; a first sparger arranged in a bottom portion of the vessel; and a second or optional more spargers provided above the first sparger to supply additional air bubbles and / or additional oxygen gas bubbles continuously to the liquid medium, wherein the second sparger is located at a position in the bioreactor or fermenter above the first sparger in a predefined distance η. A process is also described for independent management of dissolved CO2 and O2 by selecting a suitable modified gas flow rate qmod.SELECTED DRAWING: Figure 5

Description

[Technical Field]

[0001] The present invention relates to a bioreactor or fermenter for the cultivation of cells or microorganisms in suspension on an industrial scale. [Background technology]

[0002] In biopharmaceutical processes, reaching industrial scale is always of particular interest. Since maintaining culture conditions often limits the possibility of performing large-scale cultivation, increasing the degree of bioreactor volume often results in a decrease in the cellular performance of the cells to be cultivated. In addition to large-scale production, quality standards for the manufactured product must be met, while providing a reliable ability to supply the market. Therefore, reducing or eliminating the size dependency of cell culture performance is always of interest in order to achieve consistently high product titers and high product yields. As expected, higher volumetric utilization of the provided increased bioreactor or fermenter volume leads to improved productivity.

[0003] For cell culture, ensuring ideal growth conditions is crucial. In this regard, maintaining a favorable physicochemical environment, such as a desired dissolved oxygen content, culture pH value, and temperature, is crucial. However, cells are known to metabolically respond to their environment. In particular, concentration gradients can inhibit cell growth in large-scale bioreactors. Furthermore, for example, pH value has a significant impact on the surrounding medium. In a stirred bioreactor or fermenter, metabolically active cells secrete CO2, which dissolves in the surrounding liquid medium and absorbs O2 from the environment to complete cellular respiration. For example, the following reaction of CO2 in a liquid medium can be observed (pH<8.0): CO2 (gas) + H2O ⇔ H2CO3 ⇔ HCO3 - +H +

[0004] Therefore, the influx of CO into the liquid phase of a bioreactor or fermenter results in an acidic environment due to a decrease in pH. Conversely, the outgassing of CO increases the pH. Therefore, it is common practice to supply oxygen gas to the bioreactor or fermenter via a gas supply or gas supply unit called a "sparger."

[0005] Therefore, in general, to achieve high product quality and efficiency, a constant oxygen supply and a well-defined depletion of dissolved carbon dioxide by so-called (CO2-) stripping must be ensured, which should be possible to envisage in bioreactors or fermenters of any size to allow reliable scale-up of any cell or microorganism to be cultivated.

[0006] Conventional gas supply units in bioreactors or fermenters, when pure O2 is used as the gas source, initially introduce gas bubbles into a liquid phase consisting of pure O2. During the residence and rise of the bubbles in the reactor containing cells or microorganisms, O2 transfers from the gas phase to the liquid phase (reaction (1)). Conversely, CO2 formed in metabolic reactions transfers from the liquid phase to the gas phase, i.e., into the bubbles (reaction (2)). Due to the Henry's law constant (defined, for example, by Christian Sieblist et al., "Insights into Large-Scale Cell-Culture Reactors: II. Gas-Phase Mixing and CO2 Stripping," Biotechnol. J. 2011, 6, 1547-1556), the transfer of O2 and CO2 occurs at different rates.

[0007] By way of example, processes and reactions of interest that take place in a bioreactor or fermenter are shown in FIG. 1. FIG. 1 illustrates the distribution of dissolved CO in a schematic bioreactor or fermenter, e.g., having a gas supply providing pure O at the bottom near the agitator (not shown). The bubbles 10 shown in FIG. 1 are shown in three different states as they rise through the liquid: as starting simple bubbles 10.1, as bubbles 10.2 in the middle section, and as bubbles 10.3 in the upper section of the bioreactor or fermenter. Thus, starting from the gas supply at the bottom of the bioreactor, O-containing bubbles 10.1 begin to rise through the liquid medium to the surface. The central section of FIG. 1 shows the processes and reactions that occur, i.e., O gas transfers from bubbles 10.2 to the liquid phase (reaction (1)) and CO gas transfers from the liquid phase to bubbles 10.2 (reaction (2)). The Henry's law constants for O2 and CO2 gases are significantly different (Sieblist et al.; loc.cit.). When measured under the same conditions, the value for O2 is approximately 0.0013 mol / (kg*bar) and the value for CO2 is approximately 0.034 mol / (kg*bar), i.e., approximately 25 times higher, i.e., the value for CO2 is larger, which leads to accelerated diffusion. Therefore, the rate of reaction (2) is much faster than the rate of reaction (1) (indicated by the different arrow thicknesses in Figure 1). Because the Henry's constant for CO2 is higher than the corresponding value for O2, the carbon dioxide concentration in the bubbles increases more rapidly than the value for O2 decreases on their way through the reactor. As a result, the driving force for CO2 mass transfer is reduced during the bubble ascent.

[0008] In fact, bubbles 10 can supply oxygen to the liquid phase for several minutes, but their carbon dioxide uptake stops within a few seconds due to CO2 saturation. Once the bubbles are saturated with CO2, they no longer absorb CO2. This is shown in Figure 1 by bubbles 10.3. After a few seconds, bubbles 10.3 still supply oxygen to the culture but are no longer capable of absorbing CO2. Therefore, bubbles supplied to the liquid medium of a bioreactor are only active for a portion of the time due to CO2 stripping. Therefore, bubbles reach a saturated CO2 concentration after a certain rise height in the bioreactor or fermenter. Therefore, at the top of Figure 1, bubbles 10.3 no longer absorb CO2 gas, and only O2 gas transfers from bubbles 10.3 to the liquid phase. As a result, moving from the bottom to the top of Figure 1, from the bubble at state 10.1 through state 10.2 to state 10.3, the delivery of O2 into the liquid medium decreases, while the concentration of CO2 in the bubble increases until it reaches saturation. This is shown diagrammatically in the triangular arrow on the right side of Figure 1, where arrow (3) from the wide part to the arrowhead symbolizes the decreasing relative tendency for O2 delivery into the culture medium, i.e., from the gas phase to the liquid phase. Arrow (4) in Figure 1 symbolizes the increasing saturation of the bubble with CO2 from the arrowhead to the wide part, resulting in CO2 moving from the liquid phase to the gas phase, which proceeds much faster than O2 delivery. Square (5) indicates the region where bubble 10.3 is no longer able to take up dissolved CO2 from the liquid phase because it has reached its CO2 saturation concentration.

[0009] Thus, after a very short time, the bubbles 10.3 become saturated with CO2 gas and can no longer absorb more CO2 from the liquid phase, but can still release O2 gas into the liquid environment. In a bioreactor or fermenter with a permanent gas supply, a whole series of bubbles is provided that enter the liquid phase and contribute to a CO2 gradient within the liquid phase. At the bottom, the bubbles have the ability to absorb CO2, an ability that is increasingly lost as the bubbles rise to the top.

[0010] It is known that CO2 removal from the culture is a major problem, especially in large scale production bioreactors or fermenters, as stripping is mainly influenced by changes in the gas composition of the bubbles during their movement through the bioreactor or fermenter from the gas supply system towards the top end.

[0011] Therefore, as mentioned above, the management of O2- and CO2- concentrations is of particular interest in biopharmaceutical processes, especially large-scale biopharmaceutical processes. To develop strategies that allow for improved control and regulation of CO2 stripping for large-scale systems, the mass transfer performance of oxygen and carbon dioxide must be evaluated in detail. The parameters that should be observed in relation to CO2 gas are the (volumetric) carbon dioxide mass transfer coefficient k L a CO2 where k L is the transport coefficient for CO2, a is the specific interfacial area, and a = A / V L , i.e., the culture volume V L The total mass transfer cross section A per unit volume (see Christian Sieblist et al., loc.cit.) is the (volumetric) oxygen mass transfer coefficient k L a O2 The mass transfer coefficient can be based on volume, in which case it is a volumetric mass transfer coefficient.

[0012] Furthermore, on the other hand, excessive concentration levels of dissolved CO2 must be avoided, and CO2 must be removed. Insufficient CO2 stripping in large bioreactors or fermenters (i.e., those with a large height and therefore a long distance for bubbles to rise) often leads to the accumulation of dissolved CO2, which leads to high CO2 concentrations in the liquid phase and inhibits cell growth and product formation. On the other hand, carbon dioxide is required for nucleic acid synthesis, and its amount cannot be too small. Therefore, it should be noted that CO2 stripping should not have any adverse effects on the cells or microorganisms to be cultivated.

[0013] As a result, the oxygen mass transfer coefficient kL a O2 without significantly affecting the carbon dioxide mass transfer coefficient k for large-scale systems. L a CO2 Strategies need to be developed to strengthen, control and regulate the

[0014] To investigate and elucidate the interrelationship between CO2 and O2 mass transfer, various studies have been conducted to identify the effects of various operating conditions. In particular, the mixing efficiency and mass transfer performance of CO2 at laboratory and industrial scales have been investigated in detail. The results are summarized in Figures 2 and 3.

[0015] Figures 2 and 3 show different manually applied superficial gas velocities, w, at laboratory and industrial scales, respectively. 0 g Dependence of volumetric mass transfer coefficient k for carbon dioxide on volumetric agitator power input P / V at two different volumes L a CO2 Specifically, Figure 2 shows laboratory-scale experiments in an aerated stirred bioreactor or fermentor with a volume of 2 L (height in the cm range), and Figure 3 shows industrial-scale experiments in an aerated stirred bioreactor or fermentor with a volume of 12,000 L (height in the m range).

[0016] As can be expected, Figures 2 and 3 show that the volumetric mass transfer coefficient k L a CO2 In both experiments, i.e., laboratory scale and industrial scale, the superficial gas velocity w 0 g It can be seen that the volumetric mass transfer coefficient k for the industrial-scale reactor increases as the volumetric mass transfer coefficient k for the industrial-scale reactor increases. Furthermore, the mass transfer performance of carbon dioxide was found to be significantly different in the laboratory-scale compared to the industrial-scale process. However, it was unexpected that the volumetric agitator power input has a large effect on the laboratory-scale but only a minimal effect on the industrial-scale. As can be seen from Figures 2 and 3, the volumetric mass transfer coefficient k for the industrial-scale reactor L a CO2 is up to 10 times lower compared to laboratory-scale reactors.

[0017] To better illustrate the differences between laboratory-scale and industrial-scale experiments, refer to Figure 4, which shows the mass transfer coefficient k for carbon dioxide between laboratory and industrial scales. L a CO2 Figure 4 shows a comparison of the volumetric power input P / V = 21 Wm -3 Figure 1 shows the relative influence of specific power input on the volumetric carbon dioxide mass transfer coefficient compared to the volumetric carbon dioxide mass transfer coefficient measured in the laboratory and industrial scales. L a CO2 In comparison, increasing the volumetric power input P / V at industrial scale does not significantly improve mass transfer performance, but at laboratory scale it increases mass transfer performance in CO2 from 21 to 168 Wm -3 It is clear that the k value can be improved by up to 70%. Therefore, Figure 4 shows that at laboratory scale (2 L system), the k value increases with increasing agitator power. L a CO2 demonstrates a +70% increase in k L a CO2 Only a +5% increase in saturation can be observed (12,000 L system).

[0018] Therefore, in industrial-scale bioreactors or fermenters based on Figures 2 to 4, it is clear that mass transfer performance for CO2 cannot be significantly improved with increasing volumetric power input P / V, but at laboratory scale, mass transfer performance for CO2 can be improved by up to 70% (see Figure 4).

[0019] As already explained above, the different behavior of the two systems can be explained fundamentally by the residence time of the gas phase in the system. At industrial scale, equilibrium in CO2 concentration between the bubbles and the liquid is reached long before the bubbles reach the surface, whereas at laboratory scale, the residence time is too short for equilibrium to be reached. Therefore, increasing the interfacial area by increasing the agitator frequency results in a higher mass transfer coefficient at laboratory scale, but at large scale, the stronger dispersion of "dead bubbles" (i.e., bubbles with CO2 saturation) is useless.

[0020] In the case of oxygen mass transfer, further experiments have shown that equilibrium is not reached even at industrial scale. Therefore, the higher the volumetric power input, the larger the interfacial area and, therefore, the larger the volumetric mass transfer coefficient k L a O2 It is therefore concluded that the volumetric mass transfer coefficient in carbon dioxide can only be significantly improved at higher gas flow rates, but not at higher volumetric power inputs.

[0021] The difficulty of stripping carbon dioxide on an industrial scale is primarily related to the fact that the gas phase is already saturated with carbon dioxide only immediately above the submerged gas source located at or near the bottom of the bioreactor or fermenter. Therefore, the most feasible option for increasing mass transfer performance in carbon dioxide is to increase the gas flow rate. However, this also often results in an undesirable increase in the oxygen mass transfer rate, and therefore, independent control of the O2- and CO2- concentrations in the bioreactor or fermenter is not possible.

[0022] Reactors with two spargers are already known and commercially available in the prior art. For example, EP 0 099 634 A describes a reactor for multiphase contact between gas, solid, and liquid phases, comprising a cylindrical vessel, a draft tube, a conical bottom, and a gas sparger system. A gas sparger 16 is located at the lower end of the vessel in the gap between the inner wall and the periphery of the conical surface to introduce at least one gas in the form of bubbles into a continuous liquid phase in which a particulate solid phase is contained and suspended within the vessel. An auxiliary gas sparger in the form of a ring sparger 34 surrounds the draft tube and is configured to emit gas in the form of bubbles from its radially outer side into the liquid phase. EP 0 099 634 A makes no mention of the distance between the two spargers.

[0023] WO 2002 / 33048 discloses a method for cultivating microorganisms under aerobic conditions in a fermentation vessel, comprising injecting a first oxygen-containing gas into the lower part of the vessel with a non-uniform flow that causes chaotic behavior of the culture medium, and introducing a second oxygen-containing gas into the vessel, characterized in that the second oxygen-containing gas is introduced as a non-uniform flow of bubbles moving in all possible directions in the vessel, regardless of the direction of the culture medium flow, causing turbulent flow conditions at the injection site, and as a set of bubbles of non-uniform size and wide size distribution. The distance between the two spargers is not mentioned and is not important, since there are no restrictions on the inlet position of the second oxygen-containing gas stream, as outlined on page 4, lines 20-24 of the specification.

[0024] Sen Xu et al., "A practical approach in bioreactor scale-up and process transfer using a combination of constant P / V and vvm as the criterion," Biotechnology Progress, Vol. 33, No. 4, 2017, pp. 1146-1159, evaluates bioreactor scale-up as a critical step in the production of therapeutic proteins such as monoclonal antibodies (MAbs). For example, the sparged k from a range of bioreactor scales (3 to 2,000 L) with different spargers is presented. L a and k L a CO2 (CO volumetric mass transfer coefficient) is examined. In this context, single and dual sparger systems are described without any disclosure regarding their geometry. Generally, in dual sparger systems, both spargers are in approximately the same position and are not at different heights.

[0025] Furthermore, U.S. Patent No. 5,994,567, from the field of organic chemistry, is directed to a liquid-phase oxidation process using direct oxygen injection into a bubble column reactor, i.e., a first oxygen-containing gas is injected into the bottom of a bubble column reactor vessel containing an oxidizable organic liquid. A second oxygen-containing gas is further injected into the reactor at one or more points where the liquid is substantially depleted of dissolved oxygen prior to said injection. Oxygen from both the first and second oxygen-containing gases is used to oxidize the organic liquid, such as cumene or cyclohexane. Thus, while stirred-tank bioreactors or fermenters for culturing cells or microorganisms are not described, chemical reactions are described whereby stripping of CO2 produced by the cultured living cells is not important.

[0026] WO 2008 / 088371 discloses systems for containing and manipulating fluids, including systems and methods involving supported collapsible bags that can be used as reactors for carrying out chemical, biochemical, and / or biological reactions. In one embodiment, fluids contained in a vessel can be sparged, for example, to direct the fluid into the vessel's container. In some cases, the sparging can be controlled by rapidly activating or changing the degree of sparging as needed. It is noted that in some cases, multiple spargers may be used. However, the document does not mention their specific geometric arrangement. The different spargers 47 or 301 described are positioned at the same height at the bottom of the reactor, according to FIG. 1, but there is no specific teaching regarding their physicochemical effects.

[0027] Possible locations, configurations, and sizes of spargers relative to the agitator are evaluated and discussed in Sardeing et al., "Gas-Liquid Mass Transfer," Chemical Engineering Research and Design, Elsevier, Amsterdam, NL, Vol. 82, No. 9, 2004, pp. 1161-1168; Birch et al., "The Influence of Sparger Design and Location on Gas Dispersion in Stirred Vessels," Chemical Engineering Research and Design, Elsevier, Amsterdam, NL, Vol. 75, No. 5, 1997, pp. 487-496; and Rewatkar VB et al., "Role of Sparger Design on Gas Dispersion in Mechanically Agitated Gas-Liquid Contactors," Canadian Journal of Chemical Engineering, 1993, Vol. 71, No. 2, pp. 278-291. To evaluate effectiveness, only a single sparger is used. At the same time, there are two spargers in the reactor, the distance between them is not relevant and is not mentioned.

[0028] It is therefore an object of the present invention to provide a modified bioreactor or fermentor that overcomes the drawbacks of the prior art and allows for the control of carbon dioxide concentration independently of oxygen concentration within an aerated stirred bioreactor or fermentor on an industrial scale.

[0029] It is a further object to provide a method for controlling a cell culture or fermentation process by independent control of carbon dioxide and oxygen concentrations in an aerated stirred bioreactor or fermentor on an industrial scale. Summary of the Invention

[0030] Surprisingly, it has been found that the drawbacks known from the prior art can be overcome and that independent control of O2- and CO2- concentrations in industrial-scale aerated stirred bioreactors or fermenters can be achieved, in particular when a second gas source (or possibly more gas sources) is positioned a predetermined distance from the first gas source in the bioreactor.

[0031] Therefore, to overcome the aforementioned drawbacks, a modified and thereby improved bioreactor or fermentor for culturing cells or microorganisms in suspension in a liquid medium on an industrial scale is provided. A bioreactor or fermentor 100 for culturing cells or microorganisms in suspension in a liquid medium on an industrial scale, comprising: a container 102 containing a culture in a liquid medium having a determined filling height; a stirrer 120 provided in the container for stirring the liquid medium; a first sparger 150 located at the bottom 105 of the vessel 102, adapted to continuously supply gas bubbles 10, 10.1, 10.2, 10.3 to the liquid medium, the gas being selected from air and / or oxygen gas; a second sparger 160 disposed within the vessel 102 and positioned above the first sparger 150 for continuously supplying additional gas bubbles and / or additional oxygen bubbles 20, 20.1, 20.2, 20.3 to the liquid medium; Equipped with The second sparger 160 is positioned within the bioreactor or fermentor 100 a distance η above the first sparger 150, where η is at least about 0.4 m above the first sparger 150 to a maximum of about 0.5 m below the fill height of the bioreactor or fermentor 100, or From about 0.4 m above the first sparger to about 2 / 3 of the fill height of the bioreactor or fermentor 100, or From about 0.4 m above the first sparger to about ½ of the fill height of the bioreactor or fermentor 100, or The bioreactor or fermentor 100 is selected to be in the range of about 0.4 m above the first sparger to about 3.0 m above the first sparger, or about 0.4 m to about 2.5 m, or about 0.4 m to about 2.0 m, or about 0.4 m to about 1.5 m, or about 0.4 to about 1.0 m, or about 0.45 to about 0.90 m, or about 0.5 to about 0.80 m, or about 0.55 to about 0.70 m, or to be about 0.6 m.

[0032] Therefore, to improve the mass transfer performance for CO while not adversely affecting the mass transfer performance for O, according to the present invention, an additional second gas sparger is provided in the bioreactor or fermenter at a distance η, located higher than the first sparger, to achieve a much shorter residence time for the additional gas supplied compared to the gas supplied from the submerged or first sparger. The shorter residence time of the gas injected by the second sparger allows for less oxygen to be transferred to the liquid phase, while an increased amount of CO can be removed.

[0033] According to one embodiment, the second sparger may be a side sparger, ie, a sparger that provides additional bubbles near the side wall.

[0034] The prior art and embodiments of the present invention are described by way of example with reference to the accompanying drawings, which are schematic and are not intended to be drawn to scale as precise geometrical assumptions cannot be made with respect to the original size. The figures of the present disclosure are incorporated in and constitute a part of this specification and also illustrate embodiments of the present invention which are not limited to the particular embodiments described. The drawings, together with the general and detailed description, serve to explain the principles of the present disclosure. Like features are designated with like reference numerals throughout the drawings. [Brief explanation of the drawings]

[0035] [Figure 1] FIG. 1 shows a schematic diagram of the processes and reactions involved in the distribution of O and CO occurring within bubbles in a bioreactor or fermentor. [Figure 2]The volumetric mass transfer coefficient kLaCO2 for carbon dioxide in dependence of the volumetric agitator power input P / V for different superficial gas velocities w0 g at laboratory scale is shown. [Figure 3] The volumetric mass transfer coefficient kLaCO2 for carbon dioxide in dependence of the volumetric agitator power input P / V for different gas superficial velocities w0 g at industrial scale is shown. [Figure 4] Figures 2 and 3 show a comparison of the mass transfer coefficient kLaCO2 for carbon dioxide between laboratory and industrial scale based on the values ​​obtained. [Figure 5] 1 shows a schematic diagram of the processes and reactions for O2 and CO2 distribution occurring within bubbles 10 in a bioreactor or fermentor without additional gas injection (section A) and with additional gas injection (section B; bubbles 20) according to an exemplary embodiment of the present disclosure. [Figure 6a] 1 illustrates the respective locations of a first sparger 150 and a second sparger 160 within a bioreactor or fermentor according to an exemplary embodiment of the present disclosure. [Figure 6b] 1 shows the respective locations of a first sparger 150, a second sparger 160, and a third sparger 170 within a bioreactor or fermentor according to an exemplary embodiment of the present disclosure. [Figure 7] 1 illustrates types of spargers 150, 160, 170 according to exemplary embodiments of the present disclosure. [Figure 8] As described in the exemplary section (Example 1), an exemplary evaluation of kLaCO2 values ​​from saturation of cCO2=14% to cCO2=6% is shown according to an exemplary embodiment of the present disclosure. [Figure 9] FIG. 1 shows a technical diagram of a vertical cross-sectional schematic view of an industrial scale bioreactor and the installation location of an additional second sparger (160) as described in the exemplary section (Example 2) and according to an exemplary embodiment of the present disclosure. [Figure 10]The examples section (Examples 2 and 3) present a comparison of four measurements performed to determine the mass transfer coefficient for oxygen (kLaO2) in dependence of gas flow rate according to the prior art and exemplary embodiments of the present disclosure. [Figure 11] As described in the exemplary section (Example 3), a comparison of four measurements performed to determine the mass transfer coefficient for carbon dioxide (kLaCO2) in dependence of gas flow rate is presented according to the prior art as well as exemplary embodiments of the present disclosure. [Figure 12a] As explained in the exemplary section (Example 3), the determination of the influence coefficient CO2 on the mass transfer coefficient of oxygen (kLaO2) according to the present disclosure is shown. [Figure 12b] As explained in the exemplary section (Example 3), the determination of the influence coefficient CCO2 on the mass transfer coefficient of carbon dioxide (kLaCO2) according to the present disclosure is shown. [Figure 13a] As described in the exemplary section (Example 4), a comparison of four measurements performed to determine the mass transfer coefficient for oxygen (kLaO) in dependence of gas flow rate using side sparger type A or side sparger type B according to the prior art and an exemplary embodiment of the present disclosure is shown. [Figure 13b] As described in the exemplary section (Example 4), a comparison of four measurements performed to determine the mass transfer coefficient for carbon dioxide (kLaCO2) in dependence of gas flow rate using side sparger type A or side sparger type B according to the prior art and an exemplary embodiment of the present disclosure is shown. [Figure 14a] The exemplary section (Example 5) describes measurements performed to determine the mass transfer coefficient for carbon dioxide (kLaCO2) in dependence of agitator frequency and water gas flow rate. [Figure 14b] The exemplary section (Example 5) describes measurements carried out to determine the mass transfer coefficient for oxygen (kLaO2) in dependence of the agitator frequency and the gas flow rate in water. [Figure 15a]As explained in the exemplary section (Example 5), measurements were performed to determine the mass transfer coefficient for carbon dioxide (kLaCO2) in dependence on the agitator frequency and side gas flow rate. [Figure 15b] The exemplary section (Example 5) describes measurements performed to determine the mass transfer coefficient for oxygen (kLaO2) in dependence on the agitator frequency and side gas flow rate. [Figure 16a] The exemplary section (Example 6) describes the step response method, particularly the inputs to the system. [Figure 16b] The exemplary section (Example 6) describes the step response method, specifically the output of the system. [Figure 17] 1 illustrates the bubble catcher (funnel) 180 described in the exemplary section (Example 6) and used in step response measurements. [Figure 18] Typical input and output signals from the step response method described in the exemplary section (Example 6) and applied to a 12 kL aerated stirred tank reactor are shown. [Figure 19] This is explained in the exemplary section (Example 6) and shows a comparison of the step response at the bubble catcher (funnel) 180 resulting from 100% and 50% loading steps. [Figure 20] As described in the exemplary section (Example 6), the results of measurements of gas phase residence time determined by the step function response method (the "Sprungantwort-method" as described in the Examples section) in a laboratory scale bioreactor or fermentor (30 L) compared to an industrial scale bioreactor or fermentor (12,000 L) are shown. [Figure 21] As described in the Examples section (Comparative Example 1), results for viable cell density, cell viability, titer and CO2 partial pressure are shown for culturing an antibody derivative in a conventional fed-batch process in a commercially available 12,000 L bioreactor over 11 days, when the bioreactor is equipped with only one sparger. DETAILED DESCRIPTION OF THE INVENTION

[0036] A detailed description of some of the figures is provided at the end of this specification.

[0037] Terms not specifically defined herein should be given the meaning that would be given to them by one of ordinary skill in the art in light of this disclosure and the context.

[0038] A "bioreactor" is a device or apparatus in which organisms, particularly bacteria and eukaryotic cells, grow and / or synthesize useful substances, thereby consuming nutrients from the culture medium and, in the case of aerobic cells or microorganisms, O2 supplied by technical means such as spargers. In the present disclosure, a bioreactor is an industrial-scale bioreactor. A bioreactor may consist of or include a biocompatible vessel in which chemical or biochemical processes involving organisms and / or biochemically active substances derived from such organisms are carried out. Bioreactors use additional equipment, such as a stirrer, baffles, one or more spargers (e.g., as the subject of the present invention), and / or ports, that specifically enable cell culture and growth. Generally, bioreactors are in the form of a cylindrical tube with two ends, which form the top and bottom of the bioreactor. Bioreactors range in size from liters to cubic meters and are often made of stainless steel. Bioreactors according to the present disclosure are used for large-scale production.

[0039] Cultured cells, particularly eukaryotic cells such as Chinese hamster ovary (CHO) or yeast cells, are used to produce, for example, antibodies such as monoclonal antibodies and / or recombinant proteins such as recombinant proteins for therapeutic use. Alternatively, the cells may produce, for example, peptides, amino acids, fatty acids or other useful biochemical intermediates or metabolites or any other useful substance.

[0040] A "fermentor" is a device or apparatus in which microorganisms synthesize useful substances and whereby conditions suitable for the growth of the microorganisms are maintained. The above discussion regarding bioreactors applies mutatis mutandis. The fermentors of the present disclosure are used for large-scale fermentation. Known commercial products of large-scale fermentors are, for example, antibiotics, antibodies, hormones, or enzymes synthesized by such cells or microorganisms.

[0041] The produced microorganisms are useful for a variety of purposes such as wastewater treatment, the food industry for food production, the biotechnology sector for the production of drugs such as antibiotics or insulin, pest control, or biodegradation of waste, pollutants, e.g. oil pollution.

[0042] In this disclosure, the terms "industrial scale" and "large-scale" are used interchangeably and synonymously and refer to products obtained in large production volumes, often with the cost advantage of decreasing cost per unit of output as scale increases. Larger manufacturing units are expected to have lower costs per unit of output than smaller units, all other factors being equal. Industrial scale may be understood in relation to culturing cells to have a used bioreactor volume of about 2,000 L or greater. Industrial scale may be understood in relation to culturing microorganisms to have a used fermentor volume of about 1,000 L or greater. According to further embodiments, the volume of the bioreactor or fermentor used at industrial scale may be 6,000, 8,000, 10,000, 12,000, 15,000 L or greater.

[0043] An "agitator" is an object or mechanical device used for agitation, such as a magnetic agitator. Any type of agitator commonly used in cell or microbial cultivation can be used. Agitators that can be used include, for example, impellers, Rushton-Turbines, stirring paddles, and blade agitators such as pitched blade agitators.

[0044] A "sparger" is a gas source or supply device used in a bioreactor or fermentor that provides oxygen and / or gas bubbles into the liquid phase in which cells or microorganisms are cultured. Prior art bioreactors or fermentors typically have only one sparger located at or near the bottom. In this disclosure, this sparger is also referred to as the "first sparger" or "submerged sparger," and both terms are used interchangeably and synonymously.

[0045] In accordance with the present disclosure, it has been found that the provision of an additional sparger, i.e., a second sparger, positioned a predetermined distance η above the first sparger to continuously supply additional gas bubbles and / or additional oxygen bubbles to the liquid medium has various advantages for the culture process that will be apparent to those skilled in the art from the above and following descriptions.

[0046] The processes and reactions carried out in a bioreactor or fermentor according to the present disclosure are described in more detail below.

[0047] In bioreactors or fermenters with only one gas supply, typically at the bottom or lower part of the bioreactor or fermenter, bubbles entering the liquid phase have the ability to uptake CO from the cell culture medium, but this ability is increasingly lost as the bubbles rise. At the top or upper part of the bioreactor or fermenter, depending on the height as explained above, the bubbles no longer uptake dissolved CO from the liquid phase. Therefore, the CO content increases from the bottom to the top of the bioreactor or fermenter, thereby creating a CO gradient within the liquid phase.

[0048] The drawbacks associated with CO gradients and such uneven distribution of CO typically present in the liquid phase of a bioreactor or fermenter can be overcome by the present disclosure, in which a second sparger is provided a distance η above the first sparger in the liquid phase of a large-scale bioreactor or fermenter. The second sparger is provided to counteract the establishment of CO saturation of gas bubbles that can no longer absorb CO. The second sparger adds new gas bubbles to the liquid phase so that the process of passing O and CO can continue again above the first sparger.

[0049] By way of example, processes and reactions of interest that take place in the liquid phase in a large-scale bioreactor or fermenter are shown schematically in FIG. 5. In section A on the left, FIG. 5 illustrates the distribution of dissolved CO in a schematic bioreactor or fermenter with only one gas supply or sparger near the agitator (not shown). The sparger is located at or near the bottom of the bioreactor or fermenter. Standard gas treatment with only one sparger results in a clear CO gradient. After reaching a certain height, the gas bubbles 10.2 become saturated with CO. Thus, in the lower part of FIG. 5 (section A), CO stripping performance is good, while in the upper part, CO stripping is insufficient and unacceptable.

[0050] Section B on the right side of Figure 5 shows the distribution of dissolved CO2 in a large-scale schematic bioreactor or fermenter as a result of an additional second gas supply or second sparger installed above the first sparger. As can be derived from section B of Figure 5, the second sparger supplies gas bubbles 20.1 that begin to rise to the surface in the liquid medium. As shown in gas bubbles 20.2, O2 gas migrates from gas bubbles 20.2 to the liquid phase (reaction (1)), and CO2 gas migrates from the liquid phase to gas bubbles 20.2 (reaction (2)). Due to the different Henry's constants, reaction rate (2) is much faster than reaction rate (1) (shown by the different arrow thicknesses in Figure 5). Therefore, CO2 stripping in the middle and upper liquid phase in section B is similar to CO2 stripping in the lower liquid phase as shown on the left side of Figure 5 in section A. Therefore, a CO2 gradient across the rising height in the liquid medium is avoided. Cells or microorganisms present in a bioreactor or fermentor experience a more consistent environment and a much lower degree of fluctuation in the liquid medium that can be reflected in the metabolic responses of the cells or microorganisms.

[0051] The second sparger is installed above the first sparger to shorten the absolute rising height of the bubbles so as to prevent the bubbles from becoming saturated with CO2 while rising through the liquid phase. In addition, the residence time of the bubbles generated from the second sparger in the liquid phase is reduced, and the bubbles rise relatively quickly to the liquid surface because they are not dispersed by the agitator. Therefore, the CO2 stripping performance is good even in the upper part of the bioreactor or fermenter.

[0052] According to further embodiments, a third sparger and additional spargers may be present above the second sparger, etc., such that three spargers, four or more spargers are simultaneously present in the bioreactor or fermentor. The above description of the second sparger applies accordingly to the third, fourth, fifth and additional spargers.

[0053] As a result, large-scale bioreactors or fermenters are subdivided into different sub-compartments, each equipped with a sparger, in order to directly influence the CO distribution in the liquid phase. This type of "separation" of large-scale bioreactors or fermenters into smaller units leads to improved comparability and therefore predictability of industrial-scale cultures versus laboratory-scale, and vice versa, preferably resulting in higher metabolic rates, increased viability, and / or increased productivity.

[0054] The provision of a second sparger or an optional third additional sparger in a bioreactor or fermenter operating at a large production scale results in a harmonization of the physicochemical environmental conditions of the cells or microorganisms being cultured despite the large volumes used. This concept is therefore an approach that adapts the industrial scale to the laboratory scale in that the CO distribution is more uniform throughout the liquid phase.

[0055] Furthermore, the dissolved CO concentration in the liquid phase can be reduced to an appropriate concentration or level by the presence of a second sparger or an optional third sparger and further spargers. Keeping in mind that CO concentrations, especially at high concentrations, affect cell culture performance, whereby high CO concentrations inhibit the growth of aerobic cells (see David R. Gray et al., CO in large-scale and high-density CHO cell perfusion culture, Cytotechnology 1996, 22, 65-78), high CO concentrations must be avoided in the cell culture medium itself. The second sparger and further third or further spargers help prevent the occurrence of such undesirable high CO content in the liquid phase.

[0056] Additionally, additional second and third and further spargers continuously supplying additional gas bubbles and / or additional oxygen bubbles to the liquid medium also result in a more uniform O2 distribution throughout the liquid phase.

[0057] Thus, according to a further embodiment, there may be one or more further spargers located above the first, second and third spargers.

[0058] It was found that the location of the second sparger and further spargers can have a positive impact on the performance and efficiency of the culture, taking into account the management of O2- and CO2- concentrations in the liquid phase.

[0059] Assuming that the first sparger is located at or near the bottom or lower part of the bioreactor or fermenter, the second sparger is always located above the first sparger, e.g., in the middle or upper part of the bioreactor or fermenter. To further examine the location of the second sparger, there are essentially two main directions for large-scale change in the location of the second sparger within the bioreactor or fermenter. One main direction is vertical, i.e., changing the location of the sparger from the bottom to the top of the bioreactor or fermenter. That is, the second sparger can be located, for example, near the bottom or near the top of the bioreactor or fermenter, or at any distance in between. In this regard, since the gas bubbles are fed into the existing liquid phase, the fill height of the bioreactor or fermenter must be observed, not its absolute volume.

[0060] According to one embodiment of the invention, the bioreactor or fermentor comprises a fill height in the range of about 8 to about 20 m, or about 9 to about 15 m, or about 9, 5 to about 12 m, or about 10 m.

[0061] The second main direction that can be considered is the horizontal direction, i.e., the position of the sparger between the sidewall and the central axis of the bioreactor or fermenter. The central axis is an imaginary line within the bioreactor or fermenter, which is assumed to have a cylindrical shape and has equal or nearly equal spacing from the surrounding sidewall. That is, the second sparger may be located, for example, near the sidewall, or near the central axis of the bioreactor or fermenter, or at any distance therebetween.

[0062] According to further embodiments, consecutive spargers can be precisely positioned one above the other in the vertical direction. According to another embodiment, consecutive spargers may also be positioned laterally shifted relative to one another in the vertical direction.

[0063] Thus, according to the present invention, the second sparger is positioned in the bioreactor or fermenter vessel at a distance η from the first sparger, which distance η should be understood as the vertical distance, for example along the sidewall of the bioreactor or fermenter, such that the second sparger is located a distance η above the first sparger. The distance η may be in the range of at least about 0.4 m above the first sparger to a maximum of about 0.5 m below the fill height of the bioreactor or fermenter, or Within the range from about 0.4 m above the first sparger to about 2 / 3 of the fill height of the bioreactor or fermenter; or Within the range from about 0.4 m above the first sparger to about ½ of the fill height of the bioreactor or fermenter; or Within the range of about 0.4 m to about 3.0 m, or about 0.4 m to about 2.5 m, or about 0.4 m to about 2.0 m, or about 0.4 m to about 1.5 m, or Within the range of about 0.4 to about 1.0 m, or about 0.45 to about 0.90 m, or about 0.5 to about 0.80 m, or about 0.55 to about 0.70 m, or about 0.6 m above the first sparger, respectively; is selected so that

[0064] Thus, the distance η between the first and second sparger has a lower limit of about 0.4 m higher than the first sparger and an upper limit of about 0.5 m lower than the fill height of the bioreactor or fermenter. The expression "fill height," used synonymously with "liquid height," should be understood to mean the fill level of the liquid present in the bioreactor or fermenter at the start of the cultivation process, which is further defined by the surface of the liquid or the nominal volume of the liquid present. Thus, 0.5 m below the fill height of the bioreactor or fermenter is synonymous with 0.5 m below the surface of the liquid present in the bioreactor or fermenter at the start of the cultivation process.

[0065] For example, if the total filling height is 10 m, 0.5 m below the filling height is 9.5 m. Then, the distance η is selected in the range of about 0.4 m to about 9.5 m. The lower and upper limits of this range are considered to be critical values ​​of the present invention.

[0066] The presence of two spargers in the bioreactor itself increases the area of ​​the broth over which CO2 stripping occurs. If the second sparger is placed near the surface of the broth, e.g., about 0.5 m below the fill height of the bioreactor or fermentor, the second sparger can strip a downstream area, while the first sparger, placed near the bottom of the bioreactor, can strip an upstream area of ​​the broth. In essence, the entire fill height of the bioreactor or fermentor is CO2 stripped.

[0067] The distance η can be selected to be within a range from about 0.4 m above the first sparger to about ⅔ of the fill height of the bioreactor or fermenter. The expression "about ⅔ of the fill height of the bioreactor or fermenter" should be understood to mean that the second sparger is placed at a position where about ⅔ of the total fill height is achieved. For example, if the total fill height is 12 m, ⅔ of the total fill height is 8.0 m. Then, the distance η is selected in the range of about 0.4 m to about 8.0 m.

[0068] Alternatively, the distance η may be selected to be within a range from about 0.4 m above the first sparger to about ½ of the fill height of the bioreactor or fermenter. The expression "about ½ of the fill height of the bioreactor or fermenter" should be understood to mean that the second sparger is placed at a position where there is about ½ of the total fill height or about 0.5 times the liquid volume. For example, if the total fill height is 11 m, then ½ of the total fill height is 5.5 m. The distance η is then selected in the range of about 0.4 m to about 5.5 m.

[0069] The distance η may also be selected to be within the range of about 0.4 m to about 3.0 m, or about 0.4 m to about 2.5 m, or about 0.4 m to about 2.0 m, or about 0.4 m to about 1.5 m, or about 0.4 to about 1.0 m, or about 0.45 to about 0.90 m, or about 0.5 to about 0.80 m, or about 0.55 to about 0.70 m, or about 0.6 m. Thus, the distance η may be selected to be about 3.0 m, about 2.9 m, about 2.8 m, about 2.7 m, about 2.6 m, about 2.5 m, about 2.4 m, about 2.3 m, about 2.2 m, about 2.1 m, about 2.0 m, about 1.9 m, about 1.8 m, about 1.7 m, about 1.6 m, about 1.5 m, about 1.4 m, about 1.3 m, about 1.2 m, about 1.1 m, about 1.0 m, about 0.95 m, about 0.90 m, about 0.85 m, about 0.80 m, about 0.75 m, about 0.70 m, about 0.65, about 0.6 m, about 0.55, about 0.45, and about 0.4 m above the first sparger.

[0070] In a further embodiment, the distance η is within the range of at least about 0.6 m above the first sparger to a maximum of about 0.5 m below the fill height of the bioreactor or fermentor; or Within the range from about 0.6 m above the first sparger to about 2 / 3 of the fill height of the bioreactor or fermenter; or about 0.6 m above the first sparger or about 1 / 2 the fill height of the bioreactor or fermenter, or about 0.6 m to about 3.0 m above the first sparger, or about 0.6 m to about 2.5 m above the first sparger, or about 0.6 m to about 2.0 m above the first sparger, or about 0.6 m to about 1.5 m above the first sparger, or About 0.6 m to about 1.0 m above the first sparger, or about 0.6 m to about 0.90 m above the first sparger, or about 0.6 m to about 0.80 m above the first sparger, or about 0.6 m to about 0.70 m above the first sparger, or about 0.6 m above the first sparger; may be selected so that

[0071] The term "about" followed by a value should be understood to mean the value ±5% or the value ±4% or the value ±3% or the value ±2% or the value ±1%.

[0072] Of course, the location of the gas outlet opening of the sparger, i.e. the opening through which the gas bubbles enter the liquid phase, is an important criterion for determining the distance. If there are several openings in one sparger, an average value can be used to determine the appropriate distance.

[0073] The above ranges and values ​​for the distance η are the result of various experiments, calculations, and evaluations according to which the second sparger is located in the bioreactor or fermentor at a distance from the first sparger or at an upwelling height above the first sparger where the gas bubbles have reached or will soon reach CO gas phase saturation concentrations. The evaluation of CO2 gas phase saturation concentrations across the height of an industrial scale aerated stirred bioreactor or fermentor has been based on the following considerations:

[0074] According to the inventors' experiments (see Figures 2 to 4), the volumetric mass transfer coefficient k for carbon dioxide in a 12,000 L industrial-scale aerated and stirred bioreactor or fermenter is L a CO2 is the volumetric mass transfer coefficient k for carbon dioxide at laboratory scale (2 L) L a CO2 In contrast, the volumetric oxygen mass transfer coefficients k L a O2 The amounts are equivalent.

[0075] As is well known, and in contrast to oxygen mass transfer, the mass transfer of carbon dioxide from the continuous liquid phase to the gas phase is already complete after a short time. If the residence time of the gas phase in the system is longer than the time it takes for bubbles to saturate, these bubbles are no longer available for CO2 mass transfer. As a result, this effect only occurs in industrial-scale reactors, but not in laboratory-scale reactors.

[0076] To assess the period during operation when the bubbles become saturated with CO2 in an industrial-scale system (e.g., 12,000 L) and no longer contribute to CO2 stripping, the specific space-boundary interface and the CO2 mass transfer coefficient k L aCO2 needs to be taken into consideration. Therefore, first, the residence time distribution of the gas phase was evaluated as follows.

[0077] Gas phase residence time was determined by the step response method, "Sprungantwort method," which is described in detail in the experimental section. Detailed results are provided in the experimental section. The step function response method is based on the use of two different gases: one gas, such as oxygen, is used to saturate a liquid, such as water, and the other gas, such as carbon dioxide or nitrogen gas, is used to displace it and displace it from the liquid phase. Gas addition is at the bottom of the bioreactor or fermenter, and the displaced gas is measured near the top of the reactor in the liquid phase. Those skilled in the art will be familiar with this method for measuring gas phase residence time.

[0078] Thus, using the step function response method, the gas phase residence time of the bioreactor or fermentor at laboratory scale (30 L) was determined to be 5 seconds, and the gas phase residence time of the bioreactor or fermentor at industrial scale (12,000 L) was found to be 21 seconds.

[0079] Further experimental measurements and evaluations described in the experimental section then showed that the mass transfer coefficient of CO2 was 4 ± 0.68 h in a laboratory-scale bioreactor or fermenter (2 L). -1 It was identified as being.

[0080] Assuming a monodisperse distribution of bubble sizes in the 12,000 L system, d = 5 mm, the theoretical carbon dioxide profile in a single bubble can be calculated. The concentration profile of CO2 in the bubble is given by the carbon dioxide mass transfer coefficient k L a CO2 =4±0.68h -1 can be calculated assuming that it also applies to the industrial scale as follows:

number

[0081] From the experiment shown in Example 6 (shown in Figure 20), it could be estimated that after about 3.5 seconds about 95% bubble saturation could be observed.

[0082] Based on the average gas phase residence time of about 21 seconds measured in an industrial-scale bioreactor or fermenter according to the same example (FIG. 20) and the determined total distance traveled by the 3.6 m bubbles, the average bubble rise velocity can be determined to be 0.17 m / s (velocity = distance / time). As a result, after a height of about h = 0.6 m (3.5 s x 0.17 m / s), the gas phase is saturated with CO2 and therefore no further stripping of CO2 can be observed.

[0083] Because the above evaluations, measurements, and calculations involve several evaluations and assumptions, the resulting result of 0.6 m is only an approximation of the distance η, which is better represented by a range from at least about 0.4 m above the first sparger to up to about 0.5 m below the fill height of the bioreactor or fermentor.

[0084] Furthermore, experiments have shown that the presence of a second sparger located at a distance η has particular advantages for cultivating cells or microorganisms in suspension in a liquid medium on an industrial scale. The presence of a second sparger located above the first sparger at a distance η selected within the above range can be expected to have a positive effect on CO2 stripping. It can be assumed that the presence of a second sparger causes a decrease in the partial pressure of CO2 in the culture (liquid medium) of the bioreactor or fermenter. That is, when two bioreactors / fermenters are compared with each other, where both are operated under the same conditions using the same liquid medium and culturing the same cells or microorganisms, respectively, the only difference between the two bioreactors / fermenters is that one bioreactor / fermenter uses one sparger (thus reflecting the state of the art) and the other bioreactor / fermenter according to the present invention uses two spargers located at a distance η. It can then be expected that the CO partial pressure will be reduced by at least about 0.5%, or at least about 1%, or at least about 2% to about 20% in a bioreactor / fermentor with two spargers compared to one with one sparger. The extent of the reduction in CO partial pressure can be estimated based on experiments performed with one sparger (see Comparative Example 1 and FIG. 21) in conjunction with evaluation measurements in which two spargers were used (see Examples 2-6), which allowed conclusions to be drawn regarding the O and CO contents present in the bottom, middle, and top sections of the bioreactor or fermentor.

[0085] Furthermore, when a second sparger is present as in the present invention, higher product titers and higher product yields can be expected compared to bioreactors or fermenters using only one sparger. The resulting product titer or product yield can be estimated to be at least about 1%, or at least about 5%, or at least about 10% to about 30% higher than the same bioreactor or fermenter operated under the same conditions, e.g., using only one sparger. The extent of the product titer or yield can be estimated based on experiments performed with one sparger (see Comparative Example 1 and Figure 21) in conjunction with evaluation measurements in which two spargers were used (see Examples 2-6), which allowed conclusions to be drawn regarding the O2 and CO2 contents present in the lower, middle, and upper sections of the bioreactor or fermenter. It should be noted that even small improvements in processes used commercially on a large scale represent valuable technical problems to be solved. Considering the total volume of a typical large-scale bioreactor, e.g., 10,000 L or more, with millions of protein-producing cells / ml, even small improvements in yield or other industrial characteristics represent highly relevant improvements in large-scale production and must be considered significant.

[0086] Thus, according to the present invention, it has been found that the selection of the distance η in one or more of the ranges described has significant advantages and technical effects or benefits, in particular resulting in a reduction in the CO partial pressure in the culture (liquid medium) of the bioreactor and an increase in the product titer, respectively.

[0087] The presence of two spargers in a bioreactor or fermenter according to the present invention increases the total area of ​​the liquid medium through which CO2 stripping occurs. The second sparger can be positioned at a distance η, where η can be selected so that the areas through which CO2 stripping is performed by the first and second spargers overlap to some extent. On the one hand, by increasing the degree of overlap of the sparger areas, a significant improvement in the beneficial effect can be expected. On the other hand, if the two spargers are too close together, such as when the distance η is less than 0.4 m, 0.3 m, or even less, the second sparger will most likely have an increasingly smaller additional effect on the first sparger.

[0088] If the distance η is selected to be in the range from about 0.4 m above the first sparger to about ⅔ of the filling height of the bioreactor or fermenter, it can be expected that the advantageous technical effect will be more pronounced.

[0089] It can be expected that the advantageous technical effect will be even more pronounced if the distance η is selected to be in the range from about 0.4 m above the first sparger to about ½ of the filling height of the bioreactor or fermenter.

[0090] The advantageous technical effect can be expected to be particularly strong when the distance η is selected to be in the range of about 0.4 m to about 3.0 m above the first sparger, more preferably about 0.4 m to about 2.5 m above the first sparger, or about 0.4 m to about 2.0 m above the first sparger, or about 0.4 m to about 1.5 m above the first sparger, or most preferably about 0.4 to about 1.0 m above the first sparger, or about 0.45 to about 0.90 m above the first sparger, or about 0.5 to about 0.80 m above the first sparger, or about 0.55 to about 0.70 m above the first sparger, or about 0.6 m above the first sparger.

[0091] Thus, the distance η between the first and second spargers and any further spargers, if present, is each selected to be within one or more of the ranges described above.

[0092] Thus, if there are several spargers, in one embodiment the second sparger is located at a distance η 0.4 to 10 m above the first sparger. 1、2 The third sparger may be located at a distance η of 0.4 to 10 m above the second sparger. 2、3 To distinguish between different distances, the distance η between the first sparger and the second sparger may be η 1、2 and the distance between the second and third spargers is η 2、3 etc. are shown.

[0093] Thus, the distance between two successive spargers placed one above the other may be chosen to be η.

[0094] The distance η may be the same or different between all spargers present, but is always selected from the disclosed range. For better understanding, the following example is presented.

[0095] The bioreactor or fermenter comprises a vessel and has a fill height of 10 m. A first sparger is located near the bottom of the bioreactor or fermenter. A second sparger is located between the first and second spargers. 1、2 The distance η may be located within the fermentor bioreactor at a distance η, which is also denoted as 1、2 is selected to be approximately 0.6 m away from the first sparger (determined as the distance between the respective openings of both spargers where the gas bubbles enter the liquid phase). The third sparger is spaced apart from the second sparger by a distance η 2、3 In this example, η 2、3The distance between the first sparger and the third sparger is 0.6 m above the first sparger. Therefore, the third sparger is located at a distance of 2 × η above the first sparger. Therefore, the distance between the spargers is η. 1、2 and η 2、3 are equally large. The additional spargers present are η 3、4 , η 4、5 , η 5、6 , ....

[0096] However, not all distances need have the same value, but may be selected independently from the ranges disclosed above. 1、2 , η 2、3 , η 3、4 , η 4、5 , η 5、6 , ... may be selected independently of each other and may be the same or different.

[0097] According to the second main direction, the second sparger may be positioned near a side wall or near a central axis of the bioreactor or fermenter such that the gas and / or oxygen bubbles supplied from the second sparger enter the liquid phase near a side wall or near a central axis of the bioreactor or fermenter.

[0098] According to another embodiment, the second sparger may also be positioned in a position where the sparger has approximately the same distance to the sidewall and central axis of the bioreactor or fermentor at the same time.

[0099] Thus, according to one embodiment, the sparger may be a center sparger or a side sparger.

[0100] "Central sparger" should be understood in the present disclosure to mean that the central sparger is designed so that gas and / or oxygen bubbles from the sparger are delivered to the liquid phase closer to the central axis of the bioreactor or fermenter than to the side walls in the horizontal direction.

[0101] "Side sparger" in the present disclosure should be understood as a sparger designed to deliver gas and / or oxygen bubbles horizontally closer to the side wall than to the central axis of the bioreactor or fermenter.

[0102] According to one embodiment, the first sparger may be a center sparger or a side sparger.

[0103] According to another embodiment, the first sparger may be a center sparger and the second sparger may be a center sparger.

[0104] According to another embodiment, the first sparger may be a center sparger and the second sparger may be a side sparger.

[0105] According to a further embodiment, the first and second spargers may each be side spargers.

[0106] According to one embodiment, the optional third sparger may be a center sparger or a side sparger.

[0107] According to another embodiment, the first sparger may be a center sparger, the second sparger may be a center sparger, and the optional third sparger may be a center sparger.

[0108] According to another embodiment, the first sparger may be a center sparger, the second sparger may be a side sparger, and the third sparger may be a side sparger.

[0109] According to another embodiment, the first sparger may be a center sparger and all other spargers may be side spargers.

[0110] According to another embodiment, the first sparger may be a side sparger and all other spargers may be side spargers.

[0111] In one embodiment, it has been found that the advantageous technical effects disclosed herein can be significantly increased when the second sparger is a side sparger.

[0112] According to another embodiment, the second sparger may be configured so that its opening is directed downwards, i.e., so that the gas bubbles are delivered in a direction towards the bottom of the bioreactor or fermentor.

[0113] The total number of spargers in a bioreactor or fermentor can be selected, as needed, depending on the packing height present. Spargers may be present throughout the entire packing height of the bioreactor or fermentor, or only over a portion of it. The number of spargers used depends on the type of cell or microorganism selected, the dimensions of the bioreactor or fermentor, the culture conditions, etc. One skilled in the art can easily select the appropriate number of spargers for any culture system used, based on the descriptions and explanations disclosed herein.

[0114] According to a further embodiment, it has been found to be advantageous to consider the position of the agitators present in the bioreactor or fermenter with respect to the first and / or second and optional further spargers. There is no limitation on the agitators used, but any agitator chosen should have an agitator radius r positioned around a central axis A passing through the bioreactor or fermenter. s According to some experiments, the supplied bubbles are sIt has been found that it is preferable to position the first sparger at a certain distance from the central axis A of the bioreactor or fermenter so that it enters the liquid medium at a distance of 0.1 mm. In this regard, it does not matter whether the first sparger is a central sparger or a side sparger. In such cases, the sparger provides gas bubbles close to the agitator, so that maximum turbulence is achieved at the agitator blade tips. This high-energy mixing of liquid and gas is believed to be advantageous for the cultivation of cells and microorganisms.

[0115] Further experiments have shown that the second sparger and optional further spargers are positioned at a radius r of the agitator so that the bubbles being fed are outside or significantly outside the agitator motion. s It has been found that it is preferable to mount the second sparger and any further spargers at a distance from the centre greater than r. That is, the second sparger and any further spargers are preferably mounted at a distance greater than r. s The second sparger is positioned at a distance from the central axis A so as to enter the liquid phase at a distance greater than the agitator axis A. In this regard, it does not matter whether the second sparger (and any further spargers) are central or side spargers. In this case, the second sparger supplies gas bubbles spaced from the agitator so as to avoid turbulence that could deform or damage the supplied gas bubbles. High-energy mixing of liquid and gas is therefore considered detrimental to the effectiveness of the second and optional further spargers, as beneficial technical effects such as CO2 stripping performance and product yield may be adversely affected.

[0116] 6a exemplarily illustrates the locations of a first sparger 150 and a second sparger 160 within a vessel 102 of a bioreactor or fermenter 100. The bioreactor or fermenter 100 has a bottom 105 and a sidewall 110 and has a diameter D. A gas supply pipe 140, or several gas supply pipes (not shown), are located near the sidewall 110, designed and installed to form the first sparger 150 and the second sparger 160. The first sparger 150 is located at or near the bottom 105 of the bioreactor or fermenter 100 below it. In the illustrated embodiment, the first sparger 150 is a central sparger, i.e., gas bubbles enter the liquid phase closer to the central axis A than the sidewall 110 in the horizontal direction. However, as already explained, it is also conceivable that the first sparger 150 may be a side sparger.

[0117] The second sparger 160 is located above the first sparger 150. In the illustrated embodiment, the second sparger 160 is a side sparger, i.e., the gas bubbles entering the liquid phase are located closer to the side wall 110 of the bioreactor or fermenter 100 in the horizontal direction compared to their spacing relative to the central axis A. However, as already explained, it is also conceivable that the second sparger 160 may be a central sparger.

[0118] 7 shows a schematic diagram of an exemplary type of sparger that may be used as side spargers 150, 160, 170. Of course, other useful sparger types are commercially available and may be used.

[0119] In FIG. 6 a, the second sparger 160 is positioned at a height or distance η above the first sparger 150 in the bioreactor or fermentor 100. 1、2 The distance η 1、2 is within the range of at least about 0.4 m above the first sparger 150 to a maximum of about 0.5 m below the fill height of the bioreactor or fermentor 100; or within the range from about 0.4 m above the first sparger to about 2 / 3 of the fill height of the bioreactor or fermentor 100; or Within the range from about 0.4 m above the first sparger to about ½ of the fill height of the bioreactor or fermentor 100; or The second sparger 160 is selected to be within the range of about 0.4 m to about 3.0 m above the first sparger, or about 0.4 m to about 2.5 m above the first sparger, or about 0.4 m to about 2.0 m above the first sparger, or about 0.4 m to about 1.5 m above the first sparger, or about 0.4 to about 1.0 m above the first sparger, or about 0.45 to about 0.90 m above the first sparger, or about 0.5 to about 0.80 m above the first sparger, or about 0.55 to about 0.70 m above the first sparger, or about 0.6 m above the first sparger. In Figure 6a, the second sparger 160 is positioned at about half the packing height of the bioreactor or fermentor 100. As shown in Figure 6a, the second sparger 160 makes available "new" bubbles that can absorb CO from the liquid phase. As a result, dissolved CO can be better removed. In the liquid phase, a global CO gradient is largely avoided. Therefore, the environmental conditions in an industrial-scale bioreactor or fermenter are adjusted to more closely resemble those in a small-scale bioreactor or fermenter, where bubbles do not reach CO saturation concentrations upward. Therefore, stripping of growth-inhibiting dissolved CO from the culture suspension is more efficient. On the right side of Figure 6a, the two arrows (6) from the wide section to the arrowhead each symbolize the decreasing absorption capacity of CO from the liquid phase to the gas phase. The arrow (7) in Figure 6a represents the increasing saturation of the bubbles with CO from the arrowhead to the wide section. In this case, CO transitions from the liquid phase to the gas phase, and this transition occurs much faster than the delivery of O. The presence of additional spargers adjusts the environment of the cultivated cells or microorganisms more uniformly.

[0120] Further, with reference to FIG. 6a and the exemplary embodiment shown, the agitator radius r sIn this case, the agitator 120 has a shaft or axis of rotation corresponding to the central axis A. The agitator 120 is composed of three agitators R1, R2, and R3. The first agitator R1 is located above a first sparger 150, which is located at the bottom 105 or lower part of the bioreactor or fermentor 100. In addition to the first agitator R1, two further agitators R2, R3 are also provided, which are located above and below a second sparger 160. According to alternative embodiments, it is also possible to provide only the first agitator R1, or to provide both the first agitator R1 and the second agitator R2 at the same time. Embodiments with more than three agitators are also possible.

[0121] In the embodiment shown in FIG. 6a, the agitator 120 has an agitator diameter d s agitator radius r symmetrically arranged around a central axis A passing through the bioreactor or fermentor 100, resulting in s The first sparger 150 has a diameter d s or stirrer radius r s The first sparger is positioned at a distance from the central axis A (i.e., the opening of the first sparger is positioned at a distance) so that it enters the liquid phase at a distance of the following. This finding is described in Klaas Van't Riet, Review of Measuring Methods and Results in Mass Transfer in Stirred Vessels Nonviscous Gas-Liquid, Ind. Engineer Chemical Process Des. Dev., Vol. 18, No. 3, 1979, pp. 357-364, and the sparger does not need to be mounted at a distance greater than the stirrer radius from the center. The first sparger is positioned at a distance greater than the stirrer radius r from the central axis. s It is believed that the greatest turbulence is achieved at the agitator blade tips when mounted at a distance of 0.5 mm. Also, high energy mixing of liquid and gas is believed to be beneficial for culturing cells and microorganisms. Therefore, this embodiment is believed to be advantageous.

[0122] In FIG. 6a, all stirrers R1, R2, and R3 have the same dimensions, so that the stirrer radius r s and diameter d s is the same for all three agitators R1, R2, and R3 throughout the entire volume of the bioreactor or fermentor 100. Other embodiments are possible.

[0123] In the illustrated exemplary embodiment, the agitator has an agitator radius r s , and therefore agitator diameter d symmetrically arranged around a central axis A through the bioreactor or fermentor 100 s whereby the second sparger 160 has a radius r s or stirrer diameter d s The second optional additional sparger or spargers are positioned at a distance from the central axis A so as to enter the liquid phase at a distance greater than 100 m / s. If the second optional additional sparger or spargers provide gas bubbles spaced from the agitator, turbulence that could distort or damage the provided gas bubbles is avoided. Thus, high-energy mixing of liquid and gas is considered less beneficial to the effectiveness of the second and optional additional spargers, as beneficial technical effects such as CO2 stripping performance and product yield may be adversely affected.

[0124] If there are several agitators, the radius or diameter of the agitators usually refers to the agitator located next to the sparger in question.

[0125] Referring to FIG. 6b, in addition to the first sparger 150 and the second sparger 160, a third sparger 170 is provided, which is positioned at a height or distance η above the second sparger 160 within the bioreactor or fermentor 100. 2、3 The distance η 2、3 may be selected from the ranges herein defined above second sparger 160. In Figure 6b, second sparger 160 is located at about 1 / 3 of the packing height and third sparger 170 is located at about 2 / 3 of the packing height of bioreactor or fermentor 100.

[0126] The presence of a second sparger according to the present invention is expected to significantly reduce the CO partial pressure by at least about 0.5%, or at least about 1%, or at least about 2%, up to about 20%, compared to a bioreactor / fermentor having only one sparger (reflecting the state of the art). Furthermore, increased product titer and increased product yield of at least about 1%, or at least about 5%, or at least about 10%, up to about 30%, are expected to be obtained compared to a bioreactor or fermentor in which only one sparger is used. As noted above, even small improvements in processes used on an industrial scale, such as those of the present invention, represent significant improvements, for example, in terms of the possible duration of the overall process (extension from a few hours in a fed-batch process to up to one, two, or even three days), viable cell density (cell viability), or the resulting titer of the product.

[0127] Furthermore, it has been found that in bioreactors or fermenters operated on a large scale, the provision of a second sparger and possibly further spargers in addition to the first sparger already provides the possibility of establishing independent control of the O2- and CO2- concentrations within the bioreactor or fermenter.

[0128] Thus, there is provided a process for controlling and regulating the content of dissolved CO and the content of dissolved O in a liquid medium in a bioreactor or fermenter 100 for cultivating cells or microorganisms in suspension on an industrial scale, comprising a vessel 102 containing a culture in a liquid medium, the process comprising: Agitating the liquid medium; continuously supplying gas bubbles 10, 10.1, 10.2, 10.3 into the liquid medium from a first sparger 150 located at the bottom 105 of the vessel 102, the gas being selected from air and / or oxygen gas; continuously supplying gas bubbles 20, 20.1, 20.2, 20.3 into the liquid medium from a second sparger 160 disposed within the vessel 102, the gas being selected from air and / or oxygen gas, the second sparger 160 being disposed above the first sparger 150, and the second sparger 160 being a side sparger; Gas flow rate q of the submerged sparger or first sparger 150 sub and the gas flow rate q of the side sparger or second sparger 160 side Based on the corrected gas flow rate q, both of which are suitable for the culture process, mod (O2) and adjust the corrected gas flow rate q mod Selecting and adjusting (CO2); and the following equation: q mod (O2)=q sub +C O2 ×q side [1a] and q mod (CO2)=q sub +C CO2 ×q side [1b] is applied, where: q sub represents the gas flow rate in the water or first sparger 150, q side represents the gas flow rate of the side or second sparger 160; C O2 represents the influence coefficient C of volumetric oxygen mass transfer, C O2 =0.15, and C CO2 represents the influence coefficient C of volumetric carbon dioxide mass transfer, C CO2 =0.6, A process is provided.

[0129] The above process for controlling and adjusting the dissolved CO2 content and the dissolved O2 content in a liquid medium in a bioreactor or fermenter for cultivating cells or microorganisms in suspension on an industrial scale is described in detail below.

[0130] It has been recognized that more accurate predictions of the oxygen and carbon dioxide mass transfer performance of industrial-scale aerated and stirred bioreactors or fermenters can only be made on the basis of detailed studies. Therefore, numerous experiments have been performed in which only the first sparger is present, or the effect of an additional second sparger has been investigated. The oxygen mass transfer coefficient k L a O2 It has been found that the mass transfer coefficient is directly proportional to the gas flow rate when only one sparger is present; that is, doubling the gas flow rate approximately doubles the mass transfer coefficient. That is, one sparger provides sufficient oxygen mass transfer (k L a O2 This can already result in a system that is characterized by:

[0131] Furthermore, surprisingly, the presence of an additional side sparger, and therefore additional side aeration, significantly reduces the oxygen mass transfer coefficient k L a O2 Therefore, the gas flow rate in the water or first sparger in the bioreactor or fermenter alone has little or no effect on the oxygen mass transfer coefficient k L a O2 As a result, the effect of underwater aeration is considered to be dominant on oxygen mass transfer.

[0132] Furthermore, it has been observed that side injection results in a slight increase in mass transfer performance for oxygen, but a significant increase in mass transfer performance for carbon dioxide. This finding, which demonstrates another aspect of the present invention, was also completely unexpected. Indeed, the side sparger has a very strong effect on the mass transfer coefficient for carbon dioxide, which increases much more strongly with increasing side injection. As a result, the effect of side aeration on carbon dioxide mass transfer appears to be dominant.

[0133] Side injection of gas in industrial-scale aerated stirred bioreactors or fermenters affects the mass transfer coefficients of oxygen and carbon dioxide differently, allowing for independent control of oxygen and carbon dioxide concentrations. Independent control may be based on two principles: CO2 mass transfer is approximately constant if the total gas flow rate is constant, and side aeration affects CO2 and O2 mass transfer differently.

[0134] Some evaluation measurements have shown that the gas flow rate of the side gas injection cannot simply be added to the underwater gas flow rate. In practice, a "corrected" gas flow rate must be assumed. Thus, the corrected gas flow rate for carbon dioxide (q mod (CO2)) and corrected gas flow rate for oxygen (q mod (O2)) are the mass transfer coefficients (k L a O2 or k L a CO2 ) must be taken into account. Thus, the gas flow rate can be used to directly influence the CO2 mass transfer and O2 mass transfer, respectively, of the cultivation process. Experiments (illustrated by Figures 10 and 11) have shown that the "corrected" gas flow rate can be expressed as: For oxygen: q mod (O2)=q sub +C O2 ×q side [1a] where C O2 =0.15. For carbon dioxide: q mod (CO2)=q sub +C CO2 ×q side [1b] where C CO2 =0.6.

[0135] The corrected gas flow rate is the underwater aeration q sub and the side aeration q modified and weighted by the influence coefficient C side The higher the coefficient C, the greater the impact of side aeration on mass transfer performance. The coefficient C can be calculated by assuming that mass transfer is proportional to the dissolved gas flow rate. The detailed calculation of the impact coefficient C is explained and demonstrated in the Examples section (Example 3).

[0136] The results show that side injection of gas increases the mass transfer coefficient of oxygen by an impact factor of 0.15 for CO2, while the mass transfer coefficient of carbon dioxide increases by an impact factor of 0.6. CO2 That is, a slight increase in mass transfer performance for oxygen is observed, but the mass transfer performance for carbon dioxide increases significantly.

[0137] This approach allows for independent management of carbon dioxide and oxygen in industrial-scale aerated and stirred bioreactors or fermentors. In other words, the carbon dioxide influence coefficient (k L a CO2 ) is about four times higher with side injection of gas compared to underwater aeration. Therefore, the actual or corrected gas flow rates of oxygen and carbon dioxide can be selected and adjusted to have optimal conditions in the culture system.

[0138] The following typical case illustrates the above findings as follows: Gas flow rate q of the submerged or first sparger sub is chosen to have the following value: sub =120L / min. Second or side sparger gas flow rate q sideis chosen to have the following value: side =60L / min.

[0139] Next, the corrected gas flow rate for oxygen mass transfer, q mod (O2) can be calculated according to the following formula [1a]: q mod (O2)=120L / min+0.15×60L / min=129L / min

[0140] Furthermore, the corrected gas flow rate q for carbon dioxide mass transfer mod (CO2) can be calculated according to the following formula [1b]: q mod (CO2)=120L / min+0.6×60L / min=156L / min

[0141] q mod (O2) is the mass transfer coefficient k L a O2 represents the corrected total gas flow rate which is considered to be proportional to the mass transfer rate of oxygen, so those skilled in the art have a direct measure of the effect of oxygen on mass transfer. mod (CO2) is the mass transfer coefficient k L a CO2 The above equations [1a] and [1b] represent the corrected total gas flow rate that is considered to be proportional to the mass transfer rate of carbon dioxide, and those skilled in the art also have a direct measure of the effect of carbon dioxide on mass transfer. As a result, the above equations [1a] and [1b] allow for the control and adjustment of O2 content and CO2 content, respectively, based on the selected gas flow rates. Those skilled in the art will be able to determine the appropriate and desired q that is optimal for the individual culture process and has a direct impact on the culture performance of the cells or microorganisms. mod (O2) and q mod (CO2) can be easily selected and adjusted.

[0142] Thus, in this manner, it is possible to carry out independent management of carbon dioxide and oxygen in industrial scale aerated stirred bioreactors or fermentors when a second sparger in the form of a side sparger is provided. L a O2 and kL a CO2 The gas flow rates of the first and second spargers have the possibility to be adjusted and selected so that they can be controlled and adjusted for any particular process of cell or microbial cultivation.

[0143] According to one embodiment of the present invention, the gas flow rate q of the first sparger sub is the gas flow rate of the second sparger, q side may be chosen to be larger than q sub gaq side is adjusted to be larger than (q side >q sub ) can also be selected.

[0144] According to one embodiment, as previously described, the distance η between the first submerged sparger and the second side sparger 1、2 may be selected to be within the ranges defined herein. Thus, the second sparger can be placed at an elevation above the first sparger where the gas bubbles have reached or will soon reach CO gas phase saturation concentration.

[0145] In a further embodiment, a third sparger 170 and optionally one or more further spargers are provided in the bioreactor or fermentor 100 above the first sparger 150 and second sparger 160, and the distance between two consecutive spargers 150, 160 and / or 160, 170 arranged one above the other is selected to be η.

[0146] In addition to the first submerged sparger and the second side sparger, there may be additional spargers in the bioreactor or fermentor, as previously described. 1、2 , η 2、3 , η 3、4may be selected within the above ranges as previously disclosed. The additional spargers may be center spargers or side spargers, with the determination made individually for each sparger. Based on the above findings and explanations regarding the independent management of CO and O, according to one embodiment, all of the additional spargers are side spargers.

[0147] Thus, according to a further embodiment, the first sparger is a center sparger or a side sparger, and the second sparger and the optional third and further spargers are side spargers. According to a further embodiment, the first sparger is a center sparger, and the second sparger and the optional third and further spargers are side spargers.

[0148] In addition to varying the gas flow rate of the sparger used, there are other parameters and conditions that can help improve cell or microbial culture performance, such as temperature, pH or concentration of certain nutrients, agitator type, agitator speed, sparger type, sparger size, sparger shape.... Those skilled in the art are familiar with these parameters and how to modify them.

[0149] Those skilled in the art can select any sparger known from the prior art for use in cell or microbial cultivation processes. However, the type of sparger used can affect the mass transfer performance of oxygen and carbon dioxide. The number and size of the openings present, as well as the geometric shape and size of the selected sparger, can play a role. It is assumed that larger openings, higher bubble rise velocity, and therefore larger bubbles as a result of a shorter contact time between the gas and the medium, allow for stronger removal of carbon dioxide with only a weak oxygen supply. Therefore, those skilled in the art can easily select an appropriate sparger based on the average skill of the skilled artisan that is advantageous for each cultivation process.

[0150] According to one embodiment, the first, second and optional further spargers are static spargers selected from pipe-shaped spargers, for example, tube-type spargers such as open tube spargers, sinter plates, perforated slabs, ring spargers, spider-type spargers, disk-type spargers, seat-type spargers, cup-type spargers, and bushing-type spargers.

[0151] In further embodiments, the first, second and optional further spargers are the same or different spargers.

[0152] According to a further embodiment, the first, second and optional further spargers are spargers having a pipe shape, for example tubular spargers such as open tube spargers.

[0153] In a further embodiment, the first additional sparger, the second additional sparger, and the optional additional sparger are each crescent tube type spargers.

[0154] According to a further embodiment, the first further sparger, the second further sparger, and the optional further sparger are each crescent-shaped open tube spargers.

[0155] Tube-type spargers, especially open tube spargers such as crescent-shaped open tube spargers, have the advantage of good cleanability and also offer the advantage of easier cleaning in place (CIP) and sterilization in place (SIP).

[0156] Furthermore, the bioreactor or fermentor is not limited in accordance with the present disclosure. Any known aerated stirred bioreactor or fermentor can be used. Also, bubble column trickle bed reactors, loop reactors, and the like may be used.

[0157] The cells used are not limited according to the present disclosure. In one embodiment of the present disclosure, the cells may be eukaryotic cells such as mammalian cells, particularly yeast (S. cerevisiae, Pichia pastoris), Chinese hamster ovary (CHO) cells, human cells (e.g., HEK 293), or insect cells. Other cells may also be used.

[0158] Microorganisms are also not limited according to the present disclosure. In one embodiment of the present disclosure, the microorganism may be a prokaryotic cell, such as Escherichia coli or Bacillus subtilis.

[0159] The present disclosure also relates to a process for cultivating cells or microorganisms in a bioreactor or fermentor as previously described, wherein a second sparger is provided in the bioreactor or fermentor at a defined distance η to promote growth, viability, productivity and / or any other metabolic state of the cells or microorganisms to be cultivated.

[0160] The present disclosure also relates to a process for cultivating cells or microorganisms in a bioreactor or fermentor, as already described, wherein in addition to the first and second spargers, one or more further spargers are provided at a distance η as defined in the bioreactor or fermentor, respectively, to promote the growth, viability, productivity and / or any other metabolic state of the cells or microorganisms to be cultivated.

[0161] The present disclosure also provides, as previously described, a second sparger provided in a bioreactor or fermentor for culturing cells or microorganisms, wherein the second sparger is provided in the bioreactor or fermentor to promote growth, viability, productivity and / or any other metabolic state of the cells or microorganisms to be cultivated, and the second sparger is located at a position in the bioreactor or fermentor that is a distance γ above the first sparger, and η is from at least about 0.4 m above the first sparger to a maximum of about 0.5 m below the fill height of the bioreactor or fermentor, or from about 0.4 m above the first sparger to about 2 / 3 of the fill height of the bioreactor or fermentor, or From about 0.4 m above the first sparger to about ½ of the fill height of the bioreactor or fermenter; or about 0.4 m to about 3.0 m above the first sparger, or about 0.4 m to about 2.5 m above the first sparger, or about 0.4 m to about 2.0 m above the first sparger, or about 0.4 m to about 1.5 m above the first sparger, or about 0.4 to about 1.0 m above the first sparger, or about 0.45 to about 0.90 m above the first sparger, or about 0.5 to about 0.80 m above the first sparger, or about 0.55 to about 0.70 m above the first sparger, or about 0.6 m above the first sparger, For the second sparger, the sparger is selected to be:

[0162] The present disclosure also relates to a second sparger and one or more additional spargers provided in a bioreactor or fermentor for culturing cells or microorganisms, as previously described, wherein the second sparger and one or more additional spargers are provided in the bioreactor or fermentor to promote growth, viability, productivity and / or any other metabolic state of the cells or microorganisms to be cultured, and the second sparger is located at a position in the bioreactor or fermentor that is a distance η above the first sparger, where η is from at least about 0.4 m above the first sparger to a maximum of about 0.5 m below the fill height of the bioreactor or fermentor; or from about 0.4 m above the first sparger to about 2 / 3 of the fill height of the bioreactor or fermenter; or From about 0.4 m above the first sparger to about ½ of the fill height of the bioreactor or fermenter; or about 0.4 m to about 3.0 m above the first sparger, or about 0.4 m to about 2.5 m, or about 0.4 m to about 2.0 m, or about 0.4 m to about 1.5 m, or about 0.4 to about 1.0 m, or about 0.45 to about 0.90 m, or about 0.5 to about 0.80 m, or about 0.55 to about 0.70 m, or about 0.6 m above the first sparger, respectively; The second sparger and one or more further spargers are selected to be:

[0163] As a result, further injection of gas at higher positions, e.g. further side injection with larger bubbles, higher bubble rise velocity and therefore shorter contact time between gas and medium, leads to improved performance in the cell culture process.

[0164] Furthermore, according to the present invention, the oxygen mass transfer coefficient k L a O2 the volumetric carbon dioxide mass transfer coefficient k for large-scale systems without significantly affecting L a CO2 Thus, in a simplified form, a first sparger can be used to control and regulate O mass transfer, and a second sparger in the form of a side sparger and optionally one or more additional spargers can be used to control and regulate CO mass transfer. In particular, the first sparger may be used primarily to supply more O to the cultivation process, while the second and optionally one or more additional spargers may be used primarily to reduce the CO content.

[0165] Example Example 1: L a CO2 Decision Volumetric mass transfer coefficient k L a CO2A dynamic method is used to determine the volumetric mass transfer coefficient k. In the method used, the bioreactor is gasified with carbon dioxide until 15% saturation is reached. Subsequently, the desired agitator frequency n and the desired gas generation rate q are set and the decrease in carbon dioxide concentration is recorded. A plot of the recorded carbon dioxide level against the corresponding time t is used to determine the volumetric mass transfer coefficient k. L a CO2 and the following equation [2] with the saturation concentration c*, which can be explained as follows:

number

number

[0166] Example 2: First evaluation measurement To investigate the oxygen mass transfer performance of industrial-scale aerated and stirred bioreactors or fermenters, a second sparger in the form of a side sparger was designed and installed in the reactor. Measurements were performed in a 15 kL bioreactor filled with approximately 12 kL (see Table 1 below) of 0.9% (w / v) NaCl / HO. A technical diagram of the 15 kL bioreactor used and the installation location of the additional side sparger are shown in Figure 9. In these measurements, the additional side sparger 160 was installed at a position where the filling height corresponds to half the reactor volume. Fill represents the fill volume of the vessel of the bioreactor 100. That is, the side sparger 160 is Fill ) and is located at a position equivalent to 0.5 times the filling volume (1 / 2V Fill ) is in this case about 1 / 2 the fill height of the bioreactor or fermentor since the bioreactor or fermentor is cylindrical in shape.

[0167] The submersible or first sparger (not shown) used for these measurements was a side sparger in the form of a 31 x 2 mm drill, i.e., a tube-type sparger with 31 openings, each opening having a diameter of 2 mm.

[0168] The second sparger used for these measurements is also a side sparger in the form of a 15 x 3 mm drill, a tube-type sparger with 15 openings, each opening having a diameter of 3 mm.

[0169] The gas used in the submerged sparger (first sparger) and the side sparger (second sparger) is air.

[0170] Initial evaluation measurements were carried out to measure the oxygen mass transfer rate as a function of gas flow rate at an agitator frequency of n = 60 rpm. A detailed overview of the operating conditions and gasification strategies investigated is given in Table 1.

[0171] [Table 1]

[0172] Therefore, in measurements 1-4, the gas flow rate was varied while the stirrer speed was kept constant.

[0173] In Measurement 1, the gas flow rate of the submerged sparger or first sparger is set to 120 L / min. The side sparger or second sparger does not supply air to the liquid medium, i.e., the gas flow rate of the second sparger is 0 L / min. Therefore, the total or overall gas flow rate (q total ) is as follows: q sub +q side =120mL / min+0mL / min=120mL / min

[0174] In measurement 2, the gas flow rate of the submerged or first sparger was set to 100 L / min. The gas flow rate of the side or second sparger was set to 20 L / min, resulting in a total gas flow rate of 120 L / min.

[0175] In run 3, the gas flow rate of the submerged or first sparger was set to 120 L / min. The gas flow rate of the side or second sparger was set to 20 L / min, for a total gas flow rate of 140 L / min.

[0176] In run 4, the gas flow rate of the submerged or first sparger was set to 120 L / min. The gas flow rate of the side or second sparger was set to 60 L / min, for a total gas flow rate of 180 L / min.

[0177] Measurement 1, where no lateral gas treatment occurs, is used as the standard. The volumetric mass transfer coefficients k for measurements 2 to 4 are calculated relative to this standard measurement. L a O2 The volumetric mass transfer coefficient k was determined. L a O2 represents a direct measure of the oxygen mass transfer rate.

[0178] Further measurements reveal that the mass transfer coefficient kL a O2 was found to be directly proportional to the gas flow rate. When the gas flow rate is doubled, the mass transfer coefficient also nearly doubles. This finding is confirmed in Run 2, where the gas flow rate of the submerged sparger is reduced from 120 L / min to 100 L / min. In Run 1 and Run 2, the total gas flow rate is maintained at a constant level of 120 L / min, respectively, but in Run 2, the mass transfer coefficient k L a O2 The side sparger with a gas flow rate of 20 L / min has a mass transfer coefficient k L a O2 Therefore, only the gas flow rate of the submerged sparger (first sparger) has a substantial effect on k L a O2 Therefore, the effect of underwater aeration is considered to be dominant on oxygen mass transfer.

[0179] In fact, in measurements 1 to 4, the mass transfer coefficient is not significantly affected by the additional side injection of gas. Therefore, the oxygen mass transfer coefficient k L a O2 It was confirmed that there was little or no effect on

[0180] For illustrative purposes, the results obtained in Table 1 are shown in Figure 10, which shows a comparison of the four measurements performed. In Figure 10, it can be seen that reducing the submerged aeration by 16% in Measurement 2 compared to Measurement 1 has a significant effect on the O2 mass transfer coefficient, even though the total aeration rate is constant in Measurements 1 and 2. However, the improvement in side injection by 16% in Measurement 3 compared to Measurement 1 reduces the oxygen mass transfer coefficient k by 2% in Measurement 3. L a O2 This shows only an increase of 50% in run 4, and only an 8% increase in run 5.

[0181] Therefore, the first evaluation measure confirms that side injection of air has only a small effect on oxygen mass transfer performance.

[0182] Example 3: Second evaluation measurement The same measurements were performed as in the first evaluation, but the volumetric mass transfer coefficient k L a CO2 was determined (see Figure 11). From Figure 11 it can be derived that the mass transfer coefficient for carbon dioxide increases much more strongly with increasing side injection (up to 23% in run 4) compared to the mass transfer coefficient for oxygen (8%) (see Figure 10).

[0183] The first and second evaluation measurements confirm that side injection has different effects on the mass transfer coefficients of oxygen and carbon dioxide and therefore cannot simply be added to the submerged gas flow rate. As a result, side injection of gas into industrial-scale aerated stirred bioreactors or fermenters allows for independent control of oxygen and carbon dioxide concentrations, respectively.

[0184] From the above evaluation, it can be concluded that there is a "corrected" gas flow rate. A better description of the "corrected" gas flow rate can be expressed as the following equation: q mod =q sub +C×q side [1]

[0185] That is, the underwater aeration q weighted by the influence coefficient C sub and side aeration q side A corrective gas flow rate including C can be introduced. The higher the coefficient C, the greater the impact of side aeration on mass transfer performance. Coefficient C can be calculated by assuming that mass transfer is proportional to the dissolved gas flow rate. This finding is shown in Figures 12a and 12b. From Figures 12a and 12b, C O2 = 0.15, and the volumetric mass transfer coefficient of oxygen k L a O2 can be calculated, but the carbon dioxide coefficient (k L a CO2 ) is C CO2 It can be derived that =0.6 is four times higher.

[0186] As a result, additional side injection results in a slight increase in mass transfer capacity for oxygen, but a significant increase in mass transfer capacity for carbon dioxide. In fact, side injection of gas results in only a 0.15-fold increase in the mass transfer coefficient for oxygen, but can increase the mass transfer coefficient for carbon dioxide by a factor of 0.6.

[0187] Example 4: Third evaluation measurement The same measurements as in the first and second evaluation measurements were performed, but a different side sparger (second sparger) was used. Side sparger type A had a 10 x 3 mm drill, and side sparger type B had a 32 x 5 mm drill. Side sparger type B had more openings and each opening had a larger diameter compared to type A.

[0188] The results are shown in Figures 13a and 13b. Referring to Figures 13a and 13b, it can be seen that when the total gas flow rate is constant, the CO2 mass transfer is also nearly constant. Furthermore, side ventilation has a greater effect on CO2 mass transfer compared to O2 mass transfer. In particular, the volumetric mass transfer coefficient k of carbon dioxide for side sparger type B is L a CO2 increases much more strongly with increasing side injection (see Figure 13b: up to 30% in run 4) compared to the second evaluation measurement of the carbon dioxide mass transfer coefficient. This means that by choosing sparger type B, the mass transfer of CO2 can be further increased.

[0189] Therefore, it is presumed that further side injection of gas at a higher position with a greater number of larger bubbles, i.e., a higher bubble rise velocity, and therefore a shorter contact time between the gas and the medium, will result in improved performance in the cell culture process.

[0190] Example 5: Fourth evaluation measurement carbon dioxide mass transfer coefficient k L a CO2 and the mass transfer coefficient k of the oxide L a O2Further measurements were carried out to verify the dependence of on the stirrer frequency and gas flow rate, and the results are shown in Figures 14a and 14b and 15a and 15b. In the legends for Figures 14a, 14b, 15a, and 15b (provided at the end of the description), the results are summarized using plus signs ("+"). The plus signs in the legends have the following meaning: +.........Small impact +++.......Large impact

[0191] In Figures 14a and 14b, only one sparger is used, which is a submerged sparger located at the bottom of the vessel.

[0192] According to other experiments, in Figure 14a, the specific power input is proportional to the mass transfer coefficient of carbon dioxide, k L a CO2 The gas superficial flow rate has little effect on the mass transfer coefficient k of carbon dioxide. L a CO2 It was found that the effect on

[0193] Referring to FIG. 14b, oxide k L a O2 The mass transfer coefficient increases with increasing specific power input and increasing superficial gas flow rate.

[0194] In Figures 15a and 15b, the second sparger used is a side sparger. Other experiments have shown that in Figure 15a, the specific power input is proportional to the mass transfer coefficient for carbon dioxide, k, which is represented by a plus sign (+). L a CO2 The gas superficial flow rate has little effect on the mass transfer coefficient k of carbon dioxide. L a CO2 It was found that the effect on

[0195] Referring to FIG. 15b, oxide k L a O2 The mass transfer coefficient of the oxide k increases with increasing specific power input.L a O2 has only a small effect on the mass transfer coefficient of

[0196] Based on this approach, independent control of carbon dioxide and oxygen in industrial-scale aerated stirred bioreactors or fermentors is possible. Independent control may be based on two principles: CO2 mass transfer is approximately constant if the total gas flow rate is constant, and side aeration has different effects on CO2 and O2 mass transfer.

[0197] Example 6: Estimation of distance η The distance η between the first and second spargers is η 1、2 It is also called the gas phase residence time (GWP), and is determined by the CO2 saturation concentration of the bubbles and the time it takes for the bubbles to reach the CO2 saturation concentration in the bioreactor or fermenter on an industrial scale. Therefore, the gas phase residence time over the height of the reactor has been determined by a step response method ("Sprungantwort method"), the essential problems of which are explained below.

[0198] 6.1. Measurement of gas-phase residence time in an aerated stirred tank reactor using the step response method Only a few publications dealing with the determination of gas-phase residence time are available (Wachi S. and Nojima Y., Gas-Phase Dispersion in Bubble Columns, Chemical Engineering Science, Vol. 45, No. 4, pp. 901-905, 1990; Yianatos JB and Bergh LG, International Journal of Mineral Processing, 36 (1992), pp. 81-91). The gas-phase residence time distribution in two-phase flow can be determined by impulse or step response methods. However, the described impulse response methods are difficult to apply because either radioactive or toxic tracer gases are used. Furthermore, no studies on the gas-phase residence time distribution in aerated stirred-tank reactors have been published to date. Therefore, a modified measurement technique based on the step response method for determining the gas-phase residence time is used herein.

[0199] According to control theory, the behavior of a system can be determined by either the impulse response method or the step response method. The difference between both methods is the information obtained about the system. The dwell time distribution can be obtained by the impulse response method, while the dwell time itself can be obtained by the step response method. For example, the output of a system in response to a step input is shown in Figures 16a and 16b. Figures 16a and 16b illustrate the step response method, showing the input to the system (Figure 16a) and the output of the system (Figure 16b) according to Leigh, JR (2004). Control Theory 2.ed., IET control engineering series, London.

[0200] The signal measured at the output of the system is the emission of a radioactive gas tracer (Yianatos JB, Bergh LG, Duran OU, Diaz FJ, Heresi NM (1994), Measurement of Residence Time Distribution of the Gas Phase in Flotation Columns, Minerals Engineering Vol. 7, pp. 333-344) or dichlorodifluoromethane (Wachi S. and Nojima Y., loc.cit.). For practical reasons, it is often not possible to realize an input impulse using a radioactive gas tracer. Furthermore, in this case, only information about the residence time (and not the distribution) is sufficient.

[0201] Therefore, in our application with a 12 kL acrylic glass stirred tank reactor, only the gas type is changed during venting to induce a step signal in the system.

[0202] To apply the step response method, the aerated stirred tank reactor should be operated under stable process conditions, e.g., aeration, until equilibrium in the dissolved oxygen concentration is reached. During steady-state operation, the aeration is subsequently converted to pure nitrogen. The oxygen concentration is continuously measured at the inlet and outlet of the reactor. To minimize the effect of gas mixing in the reactor headspace, a funnel is installed as a "bubble catcher" above the water surface. Such a bubble catcher (funnel) 180 that minimizes the effect of the reactor headspace is exemplarily shown in Figure 17. The bubble catcher (funnel) 180 is designed to capture bubbles across its cross section and direct the collected gas to a gas sensor. As gas sensors at the reactor inlet and outlet, t response An optical oxygen sensor spot (PreSens Precision Sensing GmbH) with a very low response time of <2 seconds is used in this case. The illustrated bubble catcher (funnel) 180 in Figure 17 shows the PreSens port 185 and the off-gas 189.

[0203] Typical input signals with corresponding output signals from a step response method applied to a 12 kL aerated stirred tank reactor from the oxygen concentrations of the submerged sparger and funnel, respectively, are shown in FIG.

[0204] Assuming that the residence time scale is significantly smaller than the mass transfer time scale, the gas-phase residence time is defined as the time between the step input and the time when the oxygen concentration in the exhaust drops by more than 1%. This assumption can be verified by comparing the system responses resulting from a 100% and a 50% step input. Figure 19 shows a comparison of the step responses at the funnel resulting from a 100% and a 50% input step, respectively. With reference to Figure 19, no difference in the response signals can be detected with respect to the time when the output oxygen signal first drops.

[0205] The response step method is associated with some defects, which are not important in this case for the following reasons. When the first bubbles reach the funnel, the signal begins to drop. Depending on the bubble size distribution, this can happen very early (due to the larger bubbles), while the largest amount of small bubbles may remain in the system much longer. However, due to the use of PBS and Pluronic as solvents (phosphate buffered saline + 1 g / L Pluronic), the bubble size distribution is very narrow and the method should have acceptable accuracy. Under non-uniform flow conditions, it may happen that the main bubble plume is not captured by the funnel, in which case the residence time will be overestimated. Dissolved oxygen and nitrogen exchange between oxygen and nitrogen bubbles occurs, and further fusion and collapse may occur. This effect is considered negligible.

[0206] In summary, the step response method for determining gas phase residence time is an easily applicable method with acceptable accuracy for the described system and conditions.

[0207] The solvents used in the step response method were PBS (phosphate buffered saline) and 1.0 g / L Pluronic.

[0208] The results of the step response method are shown in Figure 20. In Figure 20, the oxygen concentration [%] is plotted against the sampling time t [s]. A laboratory-scale bioreactor or fermentor with 30 L was compared with an industrial-scale bioreactor or fermentor with 12,000 L. The agitator frequency n in [rpm] and [L h -1 The gas flow rate of the ] has been changed from laboratory scale to industrial scale to obtain equivalent power input for each system. Therefore, the agitator frequency for the industrial scale system is 300 rpm and the agitator frequency for the laboratory scale is 80 rpm. The gas flow rate for the laboratory system is 1 L min -1 and industrial-scale gas flow rate is 60 L min -1 is.

[0209] In Figure 20, two curves for laboratory scale (30 L) and two curves for industrial scale (12,000 L) are shown. Curves 1 and 2 show industrial scale measurements, while curves 3 and 4 show laboratory scale measurements. Gas concentrations are measured at the bottom (feed) of the bioreactor or fermenter and at the gas feed at the top (top) of the bioreactor or fermenter. As can be derived from the curves, the residence time t r is as follows: t r,30L =5s t r,12kL =21s.

[0210] Thus, the gas phase residence time of the laboratory scale bioreactor or fermentor was determined to be 5 seconds, and the gas phase residence time of the industrial scale bioreactor or fermentor was found to be 21 seconds.

[0211] 6.2.CO2 mass transfer coefficient k L a CO2 Rating CO2 mass transfer coefficient k L a CO2 The evaluation was carried out as follows: Specific gas boundary interface and volumetric CO2 mass transfer coefficient k in a 2L bioreactor or fermenter L a CO2 The results are summarized in Table 2 below.

[0212] [Table 2]

[0213] Based on the above measurements, the mass transfer coefficient for CO2 was 4 ± 0.68 h in a 2 L bioreactor or fermentor. -1 It was identified as being.

[0214] 6.3. Estimation of the average bubble rising velocity The monodisperse distribution of bubble sizes in a 12,000 L system is d = 5 mm, and the carbon dioxide mass transfer coefficient k L a CO2 =4±0.68h -1 Under the assumption that also applies to industrial scale, the theoretical carbon dioxide profile in a single bubble at 37°C can be calculated as follows:

number

[0215] It can be estimated that 95% saturation of the bubbles is reached after about 3.5 seconds.

[0216] Therefore, based on a measured average gas phase residence time in an industrial scale bioreactor or fermenter of approximately 21 seconds and a measured total distance traveled by the bubbles of 3.6 m, the average bubble rise velocity can be calculated as follows: Speed ​​u=distance / time: u=0.17m / s.

[0217] As a result, after a height of about h=0.6 m (3.6 m x 0.17 m / s) the gas phase is saturated with CO2 and therefore no further stripping of CO2 can be observed.

[0218] Since the above evaluations, measurements and calculations involve several estimates and assumptions, the obtained result of 0.6 m is only an approximation of the distance η, which is better represented by a range as claimed.

[0219] Comparative Example 1: Bioreactor with only one sparger A proprietary BI HEX (Boehringer-Ingelheim High Expression) CHO-DG 44-derived cell line expressing an antibody-like protein was cultured in a commercially available 12,000 L bioreactor using a conventional fed-batch process for 11 days. The bioreactor contained a Rushton and pitched blade agitator with a 2:1 H / D (height / diameter) ratio. The distance between the lower and upper impellers was 1.8 m. At nominal volume, the liquid height (=fill height) in the bioreactor was 4.2 m. The sparger was located below the lowest agitator. Growth, production, and feed media derived from the proprietary BI-HEX® platform were used in this experiment. The culture was initiated at 9,000 L and terminated at approximately 1,100 L by adding feed. Throughout the process, the culture temperature was controlled at 36.5±0.5°C, the pH was maintained in the range of 7±0.6, and the glucose concentration was maintained in the range of 0-10 g / L. Oxygenation was performed by sparging with air and oxygen. The dissolved oxygen concentration was maintained at 30%.

[0220] The results are shown in Tables A-D below and graphically in Figures 21A-D. Cells grew exponentially until day 5, after which cell number remained fairly constant. Cell viability steadily declined over the 11-day culture, ending at just below 80%. Product titers were measured starting on day 3 and increased significantly until day 11. The pCO2 profile of the cultures started at approximately 10%, decreased sharply until day 4, after which pCO2 increased again until reaching starting values ​​on day 11.

[0221] [Table 3] [Table 4] [Table 5] [Table 6]

[0222] The results of Tables A-D above are shown in Figure 21. The figure shows culture data from an exemplary 12,000 L production process according to Comparative Example 1. In (A), the normal viable cell growth curve is shown for an 11-day culture. Data is given as a percentage of the maximum cell density reached in this run. In (B), the cell viability for the culture is shown. (C) The concentration curve of the antibody derivative produced by the cells is shown. Values ​​are given as a percentage of the maximum product concentration reached during the run. In (D), the partial pressure of CO2 within the bioreactor is shown.

[0223] Example 7: Bioreactor with First and Second Sparger An experiment according to Comparative Example 1 can be performed, but as in the present invention, two spargers can be used instead of just one. The first sparger can be located below the lowest agitator and can be a center or side sparger. The second sparger can be located in the bioreactor or fermentor a distance η above the first sparger, where η is selected to be at least about 0.4 m and up to about 0.5 m above the fill height of the bioreactor or fermentor. The second sparger can be a center or side sparger.

[0224] The presence of two spargers in a bioreactor increases the area in the broth over which CO2 stripping occurs. If the second sparger is placed near the surface of the broth, e.g., about 0.5 m below the fill height of the bioreactor or fermentor, the second sparger can strip downstream areas, while the first sparger, placed near the bottom of the bioreactor, can strip upstream areas of the broth. In essence, the entire fill height of the bioreactor or fermentor is CO2 stripped.

[0225] The presence of a second sparger located above the first sparger at a distance η selected within the above range can be expected to have a significant impact on CO2 stripping, i.e., a reduction of at least about 0.5%, or at least about 1%, or at least about 2% to about 20% in the CO2 partial pressure in the bioreactor culture (broth) can be estimated compared to the embodiment of Comparative Example 1.

[0226] Additionally, higher product titers and higher product yields are expected compared to Comparative Example 1. The concentration of antibody or antibody derivative produced by the cells can be estimated to be increased by at least about 1%, or at least about 5%, or at least about 10% to about 30% compared to the embodiment of Comparative Example 1.

[0227] Even small improvements in processes used commercially on a large scale, such as in this case, of at least about 0.5% or at least about 1%, represent meaningful contributions. Even small improvements in processes such as stripping performance and yield are highly relevant improvements in large-scale production and must be considered significant.

[0228] If the second sparger is selected to be a side sparger, it can be expected that the advantageous technical effects disclosed herein, in particular the reduction in the partial pressure of CO in the culture (liquid medium) of the bioreactor and the increase in product titer, respectively, will be more pronounced.

[0229] Example 8: Variation of the distance η between the first and second spargers Experiments according to Example 7 can be performed in which the distance η between the first and second sparger is varied. The first sparger can be located below the lowest agitator, and the second sparger can be placed a distance η above the first sparger in the bioreactor or fermentor. The second sparger can be placed at a distance η above the first sparger in the bioreactor or fermentor, at about ⅔ of the fill height of the bioreactor or fermentor, about 3.0 m, about 2.9 m, about 2.8 m, about 2.7 m, about 2.6 m, about 2.5 m, about 2.4 m, about 2.3 m, about 2.2 m, about 2.1 m, about 2.0 m, about 1.9 m, about 1.8 m, or about 1.7 m. , about 1.6 m, about 1.5 m, about 1.4 m, about 1.3 m, about 1.2 m, about 1.1 m, about 1.0 m, about 0.95 m, about 0.90 m, about 0.85 m, about 0.80 m, about 0.75 m, about 0.70 m, about 0.65, about 0.6 m, about 0.55, about 0.45, and about 0.4 m above the first sparger.

[0230] The presence of two spargers in a bioreactor increases the area in the broth over which CO2 stripping occurs. The second sparger can be positioned at a distance η, which can be selected so that there is some overlap between the areas over which CO2 stripping is performed by the first and second spargers.

[0231] Therefore, the presence of a second sparger located above the first sparger at a distance η having one of the values ​​mentioned above can be expected to have a significant effect on CO2 stripping, i.e., a reduction of at least about 0.5%, or at least about 1%, or at least about 2% to about 20% in the CO2 partial pressure in the bioreactor culture (liquid medium) can be estimated compared to the embodiment of Comparative Example 1.

[0232] Additionally, higher product titers and higher product yields can be expected compared to Comparative Example 1. The concentration of antibody derivative produced by the cells can be expected to increase by at least about 1%, or at least about 5%, or at least about 10% to about 30% compared to the embodiment of Comparative Example 1.

[0233] Even small improvements in processes used commercially on a large scale, such as in this case, of at least about 0.5% or at least about 1%, represent meaningful contributions. Even small improvements in processes such as stripping performance and yield are highly relevant improvements in large-scale production and must be considered significant.

[0234] If the second sparger is selected to be a side sparger, it can be expected that the advantageous technical effects disclosed herein, in particular the reduction in the partial pressure of CO in the culture (liquid medium) of the bioreactor and the increase in product titer, respectively, will be more pronounced.

[0235] Comparative Example 2: Bioreactor with two spargers but distance η outside the claimed range Experiments according to Example 7 can be carried out in which the distance η between the first and second spargers is outside the claimed range. Specifically, the distance η is less than 0.4 m, such as 0.35 m, 0.3 m, 0.2 m, or 0.1 m. It can be expected that the technical effect of the presence of a second sparger will not be achieved, i.e., the benefits resulting from a reduction in the CO2 partial pressure in the bioreactor culture (liquid medium) and higher product titer and higher product yield will not be obtained. The positive effects of two spargers simultaneously present in the bioreactor will not occur. In fact, the performance of the bioreactor approaches that of a bioreactor with only one sparger, as described in Comparative Example 1. Therefore, the lower limit of 0.4 m can be considered a critical value. [Explanation of symbols]

[0236] 10, 10.1, 10.2, 10.3 Bubbles from the first sparger 20, 20.1, 20.2, 20.3 Bubbles from the second sparger 100 Bioreactor or Fermenter 102 Container 105 Bottom 110 Side wall 120 Stirrer 140 Gas supply pipe 150 First Sparger 160 Second Sparger 170 Third Sparger 180 Bubble Catcher (Funnel) 185 PreSens port 189 Offgas A center axis r s agitator radius d s agitator diameter d1,d2 Stirrer radius r s The distance defined by D diameter of bioreactor or fermenter R1,R2,R3,R4 Stirrer η is the distance between two spargers η1,2 Distance between the first and second spargers η 2,3 Distance between the second and third spargers η 3,4 Distance between the third and fourth spargers

[0237] Description of some drawings [Table 7] Mass transfer measurements: System: 0.9% NaCl-water / air Agitator: Rushton / Pitched Blade Capacity: 2L Temperature: 37℃ [Table 8] Mass transfer measurements: System: 0.9% NaCl-water / air Agitator: Rushton / Pitched Blade Capacity: 12000L Temperature: 37℃ Figure 4 Mass transfer measurements: System: 0.9% NaCl-water / air Agitator: Rushton / Pitched Blade Gas superficial velocity: 0.96 mm s -1 Capacity: 12000L, 2L Temperature: 37℃ [Table 9] Figure 18 Residence time measurements: Filling amount: 12m 3 Agitator: Rushton / Pitched Blade Agitator frequency: 60 rpm Gas flow rate: 60L / min Medium: DI water Temperature: T=37℃ Figure 19 Residence time measurements: Filling amount: 12m 3 Agitator: Rushton / Pitched Blade Agitator frequency: 60 rpm Gas flow rate: 60L / min Medium: DI water Temperature: T=37℃ Figure 20 -1- n=300rpm / q=1l min -1 Feed -2- n=300rpm / q=1l min -1 Top -3- n=80rpm / q=60l min -1 Feed -4- n=80rpm / q=60l min -1 Top

Claims

1. Dissolved CO in a liquid medium in a bioreactor or fermenter (100) for cultivating aerobic cells or microorganisms in suspension on an industrial scale with a volume of 2000 L or more, comprising a vessel (102) containing the culture in the liquid medium. ２ Content and dissolved O ２ A process for controlling and adjusting the content of agitating the liquid medium; continuously supplying gas bubbles (10, 10.1, 10.2, 10.3) into the liquid medium from a first sparger (150) located at the bottom (105) of the vessel (102), the gas bubbles being selected from air and / or oxygen gas; continuously supplying gas bubbles (20, 20.1, 20.2, 20.3) into the liquid medium from a second sparger (160) disposed in the vessel (102), the gas bubbles being selected from air and / or oxygen gas, the second sparger (160) being disposed above the first sparger (150), and the second sparger (160) being a side sparger; The gas flow rate q of the submerged sparger or the first sparger (150) ｓｕｂ and the gas flow rate q of the side sparger or second sparger (160) ｓｉｄｅ Based on these, the corrected gas flow rate q ｍｏｄ (O ２ ) and adjust the corrected gas flow rate q ｍｏｄ (CO ２ ) and and the following equation: q ｍｏｄ (O ２ )=q ｓｕｂ +C Ｏ２ ×q ｓｉｄｅ [1a] and q ｍｏｄ (CO ２ )=q ｓｕｂ +C ＣＯ２ ×q ｓｉｄｅ [1b] is applied, where: q ｓｕｂ represents the gas flow rate in the water or first sparger (150); q ｓｉｄｅ represents the gas flow rate of the side or second sparger (160); C Ｏ２ represents the influence coefficient C of volumetric oxygen mass transfer, C Ｏ２ = 0.15, and C ＣＯ２ represents the influence coefficient C of volumetric carbon dioxide mass transfer; C ＣＯ２ = 0.6, process.

2. q ｓｕｂ Gaq ｓｉｄｅ is adjusted to be larger than Characterized by 10. The process of claim 1.

3. the location of the second sparger (160) in the bioreactor or fermentor (100) is selected to be a distance η above the first sparger (150), where η is at least 0.4 m above the first sparger (150) to a maximum of 0.5 m below the fill height of the bioreactor or fermentor (100); or From 0.4 m above the first sparger to 2 / 3 of the fill height of the bioreactor or fermenter (100); or selected to be within the range from 0.4 m above the first sparger to ½ of the fill height of the bioreactor or fermenter (100); Characterized by 3. The process according to claim 1 or 2.

4. the location of a second sparger (160) in the bioreactor or fermentor (100) is selected to be a distance η above the first sparger (150), where η is selected to be in the range of 0.4 m to 3.0 m above the first sparger, or 0.4 m to 2.5 m above the first sparger, or 0.4 m to 2.0 m above the first sparger, or 0.4 m to 1.5 m above the first sparger; Characterized by 3. The process according to claim 1 or 2.

5. the location of a second sparger (160) in the bioreactor or fermentor (100) is selected to be a distance η above the first sparger (150), where η is selected to be in the range of 0.4 to 1.0 m above the first sparger, or 0.45 to 0.90 m above the first sparger, or 0.5 to 0.80 m above the first sparger, or 0.55 to 0.70 m above the first sparger, or 0.6 m above the first sparger; Characterized by 3. The process according to claim 1 or 2.

6. a third sparger (170) and optionally one or more further spargers are provided in the bioreactor or fermentor (100) above the first sparger (150) and the second sparger (160), the distance between two successive spargers (150, 160) (160, 170) arranged one above the other being selected to be η; Characterized by The process according to any one of claims 3 to 5.

7. the bioreactor or fermentor (100) comprises a fill height in the range of 8 to 20 m, or 9 to 15 m, or 9.5 to 12 m, or 10 m; Characterized by The process according to any one of claims 1 to 6.

8. The bioreactor or fermenter (100) comprises an agitator (120) provided in the vessel for agitating the liquid medium, The agitator (120) has an agitator radius r positioned around a central axis A passing through the bioreactor or fermentor (100). ｓ and The first sparger (150) is configured such that the supplied bubbles (10, 10.1, 10.2, 10.3) are mixed with the agitator radius r ｓ and positioned at a distance from the central axis A of the bioreactor or fermenter (100) so as to enter the liquid medium at a distance of and / or The second sparger (160) and optional further spargers are arranged such that the bubbles (20, 20.1, 20.2) fed into the second sparger (160) are spaced apart by an agitator radius r ｓ and positioned at a distance from the central axis A so as to enter the liquid phase at a distance greater than Characterized by The process according to any one of claims 1 to 7.

9. In addition to the agitator (120, R1), one or more further agitators (R2, R3, R4) are provided, and the further agitators (R2, R3, R4) are located above and / or below the second sparger (160). Characterized by 9. The process of claim 8.

10. the second sparger (160) is a side sparger; Characterized by The process according to any one of claims 1 to 9.

11. the first sparger (150) is a center or side sparger, and the second sparger (160) and optional third sparger (170) and further spargers are side spargers; Characterized by The process according to any one of claims 1 to 9.

12. the first sparger (150), the second sparger (160), and optional further spargers are static spargers selected from tube-type spargers, which are spargers with a pipe shape, open tube spargers, sinter plates, perforated slabs, ring spargers, spider-type spargers, disk-type spargers, seat-type spargers, cup-type spargers, and bushing-type spargers; Characterized by The process according to any one of claims 1 to 11.

13. 13. The process for cultivating aerobic cells or microorganisms according to any one of claims 1 to 12, wherein the bioreactor or fermentor (100) is provided with a second sparger (160) and optionally a third sparger (170) and optionally one or more further spargers to promote growth, viability, productivity and / or any other metabolic state of the aerobic cells or microorganisms to be cultivated. process.

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