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Optimization of Recombinant Enzyme Production with Pichia pastoris in an Integrated Industrial Scale-down Production Plant
11th International Symposium on Computer Applications in Biotechnology Leuven, Belgium, July 7-9, 2010
Optimization of Recombinant Enzyme Production with Pichia pastoris in an Integrated Industrial Scale-down Production Plant K. Loegering, C. Mueller, J. Fricke, H.-P. Bertelsen, U. Scheffler, R. Luttmann
Research Center of Bioprocess Engineering and Analytical Techniques HAW – Hamburg University of Applied Sciences (Tel: +49-40-42875-6357; e-mail: [email protected])
Abstract: A scale down version of a multi component production plant, including a cell breeding reactor, a protein production reactor and a disc clarifier for enzyme separation is combined to an Integrated Bioprocess. Fully automated and global observable multistage parallel process courses are developed using industrial process control systems and atline measurements for enzyme concentration and enzyme activity. Optimal production conditions were found by application of DoE - Design of Experiment. Keywords: automated enzyme production, CALB, Integrated Bioprocess, scale-down clarification, DoE, atline enzyme activity detection, Pichia pastoris
1. INTRODUCTION The methylotrophic yeast Pichia pastoris is widely used in applications for biotechnology. This well known organism is able to produce a variety of intracellular and extracellular proteins at attractive levels [Cregg et al., 1993]. An extracellular lipase CALB was expressed, which has many applications in the area of white biotechnology. It is one of the most widely used biocatalysts in organic synthesis [Anderson et al., 1998], it is stable in organic solvents [Degn et al., 2001] and active towards a wide range of substrates [Rotticci et al., 1998]. CALB was used exemplarily as an enzyme for an industrial application with a scale-down production plant. This plant consists of two bioreactors and a disc clarifier. It includes a high optimization potential in the improvement of the three single operations but also in the combination of sequential processing steps. 1.1 Strain As a production host system a Pichia pastoris wildtype strain X-33 was transformed with a pPICZ-A vector containing the information for a Lipase B from Candida antarctica (CALB). As shown in Fig. 1 both AOX genes remained intact, therefore the strain is methanol utilization positive (Mut+).
The vector contains an -factor from Saccharomyces cerevisiae for the secretion of the enzyme into the supernatant. The gene information for a Zeocin® resistance was used for the selection of mutants. The expressed enzyme consists of 317 amino acids and has a molecular weight of 33 kDa [Uppenberg et al., 1994]. 1.2 Materials Stock cultures of Pichia pastoris strain X-33 CALB were stored at -80°C in 15% (v/v) Glycerol and 0.9% (w/w) NaCl. Cultivation was done with FM22 minimal media (30 gl-1 glycerol, 25.7 gl-1 potassium dihydrogen phosphate, 5 gl-1 ammonium sulfate, 8.6 gl-1 potassium sulfate, 1.4 gl-1 calcium sulfate dihydrate, 16.4 gl-1 magnesium sulfate heptahydrate, 5.88 gl-1 sodium citrate dihydrate, 8 mll-1 vitamin solution and 4 mll-1 trace elements solution). The vitamin solution contained 0.2 gl-1 biotin, the trace elements solution contained 2 gl-1 cupper (II)-sulfate pentahydrate, 80 mgl-1 sodium iodide, 3 gl-1 manganese sulfate hydrate, 0.2 gl-1 disodium molybdenum dihydrate, 20 mgl-1 boric acid, 0.5 gl-1 calcium sulfate dihydrate, 0.5 gl-1 cobalt (II)-chloride hexahydrate, 7 gl-1 zinc sulfate heptahydrate, 22 gl-1 iron sulfate heptahydrate and 1 mll-1 sulfuric acid. For media preparation FM22 media was autoclaved at 121°C. The vitamin and trace elements solution was filter sterilized and aseptically combined with the media. The pH was adjusted to 5.0 with 12.5% ammonium hydroxide inside the bioreactor. 2. DEVELOPMENT OF AN INTEGRATED BIOPROCESS 2.1 Objectives
Fig. 1: Scheme of the Pichia pastoris transformation
978-3-902661-70-8/10/$20.00 © 2010 IFAC
A fully automated self optimizing production plant in a scaledown version is under development within the research cluster BIOCATALYSIS 2021. The first steps to archive this objective are presented here.
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2.2 Plant overview and process strategy
2.3 Cell breeding cycles
In Fig. 2 the concept of the three stage Integrated Bioprocess is shown.
At the beginning of each breeding cycle the residual broth volume is 0.7 l. This is diluted with 8.3 l of the glycerol feed stock which results in an approximate cell density of 5 gl-1, shown in Fig. 3. For a period of 12 h an unlimited growth on glycerol at a temperature L of 30°C follows before the Pichia cells are induced with methanol at a controlled concentration cS2M of 0.5 gl-1 for another 12 hours. The induction in the second part of the cycle is done after an automatic batch end detection. During induction the temperature is set to 20°C to reduce protease activity. At the end of each cycle a volume of 8.5 l with a cell density of approximately 30 gl-1 is ready for a transfer into the production bioreactor. Before discussing the next step of enzyme production in the integrated bioprocess, the optimization of the expression parameters is explained.
Fig. 2: Concept of a three stage Integrated Bioprocess The production plant consists of a 10 l cell breeding bioreactor BIOSTAT® ED10 (Sartorius, Germany), a 30 l production bioreactor BIOSTAT® C30 (Sartorius, Germany) and a semicontinuously disk clarifier SC1 (GEA Westfalia Separator, Germany). The cell breeding bioreactor is supplied with fresh media containing glycerol at a concentration of 43.8 gl-1. The production bioreactor is supplied with FM22 minimal media without a carbon source. Both reactors are fed with methanol during induction of enzyme expression. The three stages cell breeding, production and clarification are performed in parallel cycles, where the second and the third stage build on their previous stage. Therefore the cell breeding is 24 h in advance of the production and the clarification completes each production cycle.
3. OPTIMIZATION OF ENZYME PRODUCTION The optimization of target protein expression was performed with the methods of Design of Experiments (DoE). The analysis and evaluation of produced data was done with the software program Modde® [Eriksson et al., 2008]. 3.1 Screening factors Four cultivation parameters were tested in a fractional factorial 24-1 design. Initially a cube was generated as search domain for three factors. Therefore eight corner experiments had to be performed plus three experiments in the middle of the parameter ranges (center points). From this design one three-factor interaction term results besides three linear and three two-factor interaction terms. The three-factor interaction has a low significance and was used as a fourth linear term to screen four factors instead of three factors with a now cubic fractional factorial design, which is shown in Fig. 4.
Fig. 3: Four cycles of a fully automated cell breeding process NSt: agitation speed, pO2: dissolved oxygen tension, L: cultivation temperature (-), cS2M: methanol concentration (•), cXL: cell density (○)
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Fig. 4: Search area of 24-1 fractional factorial design
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volumetric oxygen uptake rate QO2, the volumetric carbon dioxide evolution rate QCO2 and the respiratory quotient RQ to monitor the respiration activity of the cells.
3.2 Automation structure of a DoE-research bioreactor
3.3 Performance of DoE design From previous experiments was known that the cultivation temperature L has an influence on the productivity because protease activity is inhibited at lower temperatures. Therefore an optimal temperature had to be found for lowest protease activity. The pH during cultivation could also change the protease activity and was under investigation as well. The influence of the inducer concentration cS2M (methanol) and of glycerol addition in an exponential feed profile with a set point µS1w had to be determined as two additional parameters. The search domain of optimal cultivation conditions is shown in tab. 1. The experiments were performed as sequential cycles in a repeated methanol fed batch with FM22 minimal media. Fig. 5: Automation structure of a DoE-research bioreactor A research bioreactor is build up with various measurement and control units as shown in Fig. 5. Besides the standard control task of maintaining a constant temperature, the pH is titrated with 12.5% ammonium hydroxide or 2 M phosphoric acid and the dissolved oxygen tension (pO2) is controlled to a set point of 25% in a cascade with agitation speed or with oxygen blending. The inducer concentration is measured inline with a diffusion probe (Kempe, Germany) and manipulated with a peristaltic methanol pump (Watson Marlow, UK). The turbidity (optek-Danulat, Germany) as well as the broth capacitance (Aber, UK) are measured inline to estimate the cell density and cell viability. An online off gas measurement of O2 and CO2 (BlueSens, Germany) provides the
After an automated harvest procedure the residual cells were diluted with fresh media and used as initial condition to grow on methanol in a new sequence. Each cycle started with a volume of 3.7 l and a cell density of 15 gl-1. The screening parameters were changed every sequence. Center point experiments were performed in the beginning, in the middle and at the end of the cultivation. A time course of these experiments is shown in Fig. 6. The total protein concentration cPtotM was determined with a Bradford Kit (BioRad, Germany). The target protein concentration cP1M was quantified from a SDS-PAGE using a fluorescent staining and a Multi-Imager (BioRad, Germany). During optimization of enzyme production a target protein activity measurement was not available.
Table 1: Search domain for parameter screening low
center point
high
L 18°C
22°C
26°C
pH 4.5
5
5.5
cS2M 0.5 gl-1
1.75 gl-1
3 gl-1
µS1 0 h-1
0.025 h-1
0.05 h-1 Fig. 6: Fully automated screening steps in a fractional factorial design
µS1w: set point of exponential glycerol feed, cP1M: target protein concentration (□), cPtotM: total protein concentration (◊), PRDj: target protein productivity
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3.4 Result of DoE screening The enzyme concentration cP1M and the cell density cXL were determined at the beginning (tj) and at the end (tj+1) of each experiment. The performance index of this cycle j is the mean productivity PRD, VM (t j1 ) cP1M (t j1 ) VM (t j ) c P1M (t j ) , PRD j (1) VL (t j1 ) t j1 t j which is related to the liquid volume VL at the end and the duration of this cycle. For a correct calculation of the mass of secreted protein, the volume of the media VM, c (t ) VM (t k ) 1 Z / X XL k VL (t k ) , k = j, j+1 Z
(2)
has to be calculated with the cell density cXL, using the weight ratio Z/X between wet and dry cells and the density Z of the wet cells. The experimental determined productivity with (1) revealed two significant and two non significant parameters with a response variability R2 of 0.8 and a predictive ability Q2 of 0.7. A significant influence on the expression of CALB was found in the inducer concentration cS2M (methanol) and in the cultivation temperature L. An exponential glycerol feed with its set point µS1w and a variation of the pH had no significant influence on the protein expression. In Fig. 7 the resulting response contour plot of the productivity is shown. Only the two significant parameters cS2M and L were plotted, whereas the expression rates are represented in colors from red (high productivity) to blue (low productivity). A high enzyme expression rate was found in the lower left corner (red) at a low temperature and a low methanol concentration.
A cultivation temperature below 18°C is not possible to maintain inside the DoE-research bioreactor. A methanol concentration of 0.5 gl-1 is the limit of quantification. Therefore this corner was taken as the optimal conditions for a high productivity. Using this knowledge the enzyme production was carried out as described below. 4. FULLY AUTOMATED PRODUCTION Each enzyme production cycle began with a transfer of 225 g cells from the cell breeding reactor. The broth was diluted in fresh media to a volume of 15 l resulting in a start cell density of 15 gl-1. The needed transfer volume was calculated with the cell density of the cell breeding reactor, estimated from the turbidity signal. A production cycle last 24 hours and was carried out at a cultivation temperature of 18°C, pH 5 and a controlled methanol concentration of 0.5 gl-1 for high expression rates as described above. No bi-substrate feed with glycerol was performed. In Fig. 8 nine production cycles are presented, where cycle 2.9 last 48 hours and shows no decrease of productivity in an extended expression time. Production cycle 2.5 follows a bad cell breeding cycle 1.5 with accidentally no pH control. This shows the dependency on a good cell breeding for high expression rates in the production. On the other hand the process strategy offers a complete regeneration of the product expression, as seen in the protein concentration chart in cycle 2.6. A new 12 hour batch phase with unlimited growth on glycerol in the cell breeding step probably causes this process recovery after a cycle with poor conditions.
Fig. 7: Response contour plot of parameter screening
Fig 8: Protein production cycles in the fully automated Integrated Bioprocess AP1M: UV absorption of enzyme concentration (∆), cS2M: inducer concentration (•), AP1actM: enzyme activity against tributyrin (□), pO2: dissolved oxygen tension
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4.2 Process automation with S88 The process automation was done with a S88 programming in the SCADA system MFCS/win 3.0 (Sartorius, Germany). A part of the block structure for the automation is shown in Fig. 9. These are the steps for an automated harvest of the production bioreactor into the disc clarifier, the refresh steps to fill fresh media into the bioreactor and the cell transfer steps from the cell breeding into the production bioreactor. This can be initialized via a manually changed variable value or by batch time.
100 mM sodium hydroxide (TitriPak, Merck, Darmstadt). One unit of lipase activity was defined as the formation of one µmol butyric acid per minute. 5.2 Atline target protein concentration The target protein concentration was measured every 35 minutes via HPLC (LaChrom Elite, VWR, Germany) at 20°C with a modified method from Trodler et al. A buffer exchange of the sample was performed with a size exclusion chromatography step using Sephadex G25 superfine in a 5 ml column (HiTrap Desalting, GE). CALB was bound (20 mM sodium citrate, pH 3) to a strong cation exchanger in a 1 ml column (HiTrap SP FF, GE) and eluted (30 mM sodium acetate, pH 5.5) with increasing the pH of the buffer. The absorption was measured at 280 nm.
Fig. 10: Atline measurements for lipase activity (up right) and concentration of the target protein (down right) 6. INTEGRATED PROCESSING Fig. 9: S88 programming structure for the initialization of a new production cycle
6.1 Cell harvest with disc clarifier
5. ATLINE PRODUCT OBSERVATION
The whole culture of the production bioreactor is clarified in the last stage via a semi-continuously disk clarifier SC1 with optimal parameters of 12500 rpm rotational speed, 40 lh-1 inlet flow and 3.4 bar outlet pressure.
As shown in Fig. 10, a cell free sample flow from the production bioreactor was carried out at 15 mlh-1 with a peristaltic pump through a filtration probe with a pore size of 0.2 µm (ESIP, Trace, Germany). The supernatant was collected in a vial, which level was maintained at 6 ml with an outlet flow. An automatic titrator dispersed the sample to the sample loop of a HPLC device and into the reaction chamber of the titrator after an optional dilution. 5.1 Atline activity measurement against tributyrin The hydrolytic activity of the enzyme against tributyrin was measured atline using an automatic titrator in pH-stat mode (ProcessLab 875, Metrohm, Germany). The reaction was carried out at 36°C and initiated by adding 100 µl of the sample to the substrate solution (5% tributyrin, 2% gum arabicum). The pH was controlled to 7.0 by titration with
For the optimization of the clarification the inlet flow and the outlet pressure were under research (data not shown). With a solid holding space of 0.7 l and an inlet flow of 40 lh-1 a cell discharge was done every six minutes followed by an approximately six minute recirculation until the cell separation was stabilized again. The clarification was satisfying because a reduction of solid content with a factor of 500 was reached, represented by the ratio between the turbidity of the inlet and outlet flow. In Fig. 11 one cycle of the whole Integrated Bioprocess is shown exemplarily. As presented in cycle 3.3 the enzyme activity AP1actC was measured in the clarifier outlet and in samples from the recirculation. No significant loss of product quality during the first down-stream step was determined.
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Fig. 11: One cycle of the whole Integrated Bioprocess
Ein, Eout: turbidity in clarifier inlet and outlet, AP1M: UV absorption of enzyme concentration (∆), AP1actK: activity of enzyme against tributyrin (□), Fout: outlet flow of the clarifier
7. OUTLOOK A further reduction of impurities from the supernatant such as cell debris or DNA could be done with a micro filtration. SDS-PAGE analysis showed already a high purity of the supernatant after clarification. For an industrial application of CALB no further purification is necessary. The next steps will be an ultra filtration to concentrate the supernatant and a lyophilization. With the opportunity to monitor the product quantity as well as the product activity in the production reactor and an additional application of these measurements in the down stream steps a fully automated plant offers a high potential to develop a self optimizing integrated production. This will be done with online applications of heuristic parameter search algorithms [Nelder and Mead, 1965] or genetic algorithm for parameter identification [Takors et al., 2001] in the future.
Eriksson, L., Johansson, E., Kettaneh-Wold, N., Wikström, C., Wold, S. (2008): Design of Experiments – Principles and Applications, MKS Umetrics AB, Umea, Sweden. Nelder, J.A., Mead, R. (1965): A simplex method for function minimization. Computer Journal (7), 308-313. Rotticci, D., Haeffner, C., Orrenius, C., Norin, T., Hult, K. (1998): Molecular recognition of sec-alcohol enantiomers by Candida antarctica lipase B. J. Mol. Catal. B. 5, 267-272. Takors, R., Weuster-Botz, D., Wiechert W., Wandrey C. (2001): A Model Discrimination Approach for Data Analysis and Experimental Design. In Hofman, M. and Thonart, P. (ed): Engineering and Manufacturing for Biotechnology, Vol. IV, 111-128. Kluver Academic.
REFERENCES
Trodler, P., Nieveler, J., Rusnak, M., Schmid, R.D., Pleiss, J. (2008): Rational design of a new one-step purification strategy for C. antarctica lipase B by ion-exchange chromatography. J Chromatogr A 1179, 161-167.
Anderson, E.M., Larsson, K.M., Kirk, O. (1998): One biocatalysis - many applications: The use of Candida antarctica B-lipase in organic synthesis. Biocatal Biotransform 16 (3), 181-204.
Uppenberg, J., Hansen, M.T., Patkar, S., Jones, T.A. (1994): The sequence, crystal structure determination and refinement of two crystal forms of lipase B from Candida antarctica. Structure 2 (4), 293-308.
Cregg, J.M., Vedvick, T.S., Raschke, W.C. (1993): Recent advances in the expression of foreign genes in Pichia pastoris. Bio/Technology 11, 905-910.
ACKNOWLEDGEMENTS
Degn, P., Zimmermann W. (2001): Optimization of carbohydrate fatty acid ester synthesis in organic media by a lipase from Candida antarctica. Biotechnol. Bioeng. 74, 483-491.
The authors thank the BMBF - German Federal Ministry of Education and Research for the financial support, Project BIOKATALYSE2021 - P13: FKZ 0315167, the Institute of Technical Biochemistry at the University of Stuttgart for providing the X-33 Pichia pastoris strain and industrial partners for their support.
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Optimization of Recombinant Enzyme Production with Pichia pastoris in an Integrated Industrial Scale-down Production Plant K. Loegering, C. Mueller, J. Fricke, H.-P. Bertelsen, U. Scheffler, R. Luttmann
Research Center of Bioprocess Engineering and Analytical Techniques HAW – Hamburg University of Applied Sciences (Tel: +49-40-42875-6357; e-mail: [email protected])
Abstract: A scale down version of a multi component production plant, including a cell breeding reactor, a protein production reactor and a disc clarifier for enzyme separation is combined to an Integrated Bioprocess. Fully automated and global observable multistage parallel process courses are developed using industrial process control systems and atline measurements for enzyme concentration and enzyme activity. Optimal production conditions were found by application of DoE - Design of Experiment. Keywords: automated enzyme production, CALB, Integrated Bioprocess, scale-down clarification, DoE, atline enzyme activity detection, Pichia pastoris
1. INTRODUCTION The methylotrophic yeast Pichia pastoris is widely used in applications for biotechnology. This well known organism is able to produce a variety of intracellular and extracellular proteins at attractive levels [Cregg et al., 1993]. An extracellular lipase CALB was expressed, which has many applications in the area of white biotechnology. It is one of the most widely used biocatalysts in organic synthesis [Anderson et al., 1998], it is stable in organic solvents [Degn et al., 2001] and active towards a wide range of substrates [Rotticci et al., 1998]. CALB was used exemplarily as an enzyme for an industrial application with a scale-down production plant. This plant consists of two bioreactors and a disc clarifier. It includes a high optimization potential in the improvement of the three single operations but also in the combination of sequential processing steps. 1.1 Strain As a production host system a Pichia pastoris wildtype strain X-33 was transformed with a pPICZ-A vector containing the information for a Lipase B from Candida antarctica (CALB). As shown in Fig. 1 both AOX genes remained intact, therefore the strain is methanol utilization positive (Mut+).
The vector contains an -factor from Saccharomyces cerevisiae for the secretion of the enzyme into the supernatant. The gene information for a Zeocin® resistance was used for the selection of mutants. The expressed enzyme consists of 317 amino acids and has a molecular weight of 33 kDa [Uppenberg et al., 1994]. 1.2 Materials Stock cultures of Pichia pastoris strain X-33 CALB were stored at -80°C in 15% (v/v) Glycerol and 0.9% (w/w) NaCl. Cultivation was done with FM22 minimal media (30 gl-1 glycerol, 25.7 gl-1 potassium dihydrogen phosphate, 5 gl-1 ammonium sulfate, 8.6 gl-1 potassium sulfate, 1.4 gl-1 calcium sulfate dihydrate, 16.4 gl-1 magnesium sulfate heptahydrate, 5.88 gl-1 sodium citrate dihydrate, 8 mll-1 vitamin solution and 4 mll-1 trace elements solution). The vitamin solution contained 0.2 gl-1 biotin, the trace elements solution contained 2 gl-1 cupper (II)-sulfate pentahydrate, 80 mgl-1 sodium iodide, 3 gl-1 manganese sulfate hydrate, 0.2 gl-1 disodium molybdenum dihydrate, 20 mgl-1 boric acid, 0.5 gl-1 calcium sulfate dihydrate, 0.5 gl-1 cobalt (II)-chloride hexahydrate, 7 gl-1 zinc sulfate heptahydrate, 22 gl-1 iron sulfate heptahydrate and 1 mll-1 sulfuric acid. For media preparation FM22 media was autoclaved at 121°C. The vitamin and trace elements solution was filter sterilized and aseptically combined with the media. The pH was adjusted to 5.0 with 12.5% ammonium hydroxide inside the bioreactor. 2. DEVELOPMENT OF AN INTEGRATED BIOPROCESS 2.1 Objectives
Fig. 1: Scheme of the Pichia pastoris transformation
978-3-902661-70-8/10/$20.00 © 2010 IFAC
A fully automated self optimizing production plant in a scaledown version is under development within the research cluster BIOCATALYSIS 2021. The first steps to archive this objective are presented here.
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2.2 Plant overview and process strategy
2.3 Cell breeding cycles
In Fig. 2 the concept of the three stage Integrated Bioprocess is shown.
At the beginning of each breeding cycle the residual broth volume is 0.7 l. This is diluted with 8.3 l of the glycerol feed stock which results in an approximate cell density of 5 gl-1, shown in Fig. 3. For a period of 12 h an unlimited growth on glycerol at a temperature L of 30°C follows before the Pichia cells are induced with methanol at a controlled concentration cS2M of 0.5 gl-1 for another 12 hours. The induction in the second part of the cycle is done after an automatic batch end detection. During induction the temperature is set to 20°C to reduce protease activity. At the end of each cycle a volume of 8.5 l with a cell density of approximately 30 gl-1 is ready for a transfer into the production bioreactor. Before discussing the next step of enzyme production in the integrated bioprocess, the optimization of the expression parameters is explained.
Fig. 2: Concept of a three stage Integrated Bioprocess The production plant consists of a 10 l cell breeding bioreactor BIOSTAT® ED10 (Sartorius, Germany), a 30 l production bioreactor BIOSTAT® C30 (Sartorius, Germany) and a semicontinuously disk clarifier SC1 (GEA Westfalia Separator, Germany). The cell breeding bioreactor is supplied with fresh media containing glycerol at a concentration of 43.8 gl-1. The production bioreactor is supplied with FM22 minimal media without a carbon source. Both reactors are fed with methanol during induction of enzyme expression. The three stages cell breeding, production and clarification are performed in parallel cycles, where the second and the third stage build on their previous stage. Therefore the cell breeding is 24 h in advance of the production and the clarification completes each production cycle.
3. OPTIMIZATION OF ENZYME PRODUCTION The optimization of target protein expression was performed with the methods of Design of Experiments (DoE). The analysis and evaluation of produced data was done with the software program Modde® [Eriksson et al., 2008]. 3.1 Screening factors Four cultivation parameters were tested in a fractional factorial 24-1 design. Initially a cube was generated as search domain for three factors. Therefore eight corner experiments had to be performed plus three experiments in the middle of the parameter ranges (center points). From this design one three-factor interaction term results besides three linear and three two-factor interaction terms. The three-factor interaction has a low significance and was used as a fourth linear term to screen four factors instead of three factors with a now cubic fractional factorial design, which is shown in Fig. 4.
Fig. 3: Four cycles of a fully automated cell breeding process NSt: agitation speed, pO2: dissolved oxygen tension, L: cultivation temperature (-), cS2M: methanol concentration (•), cXL: cell density (○)
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Fig. 4: Search area of 24-1 fractional factorial design
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volumetric oxygen uptake rate QO2, the volumetric carbon dioxide evolution rate QCO2 and the respiratory quotient RQ to monitor the respiration activity of the cells.
3.2 Automation structure of a DoE-research bioreactor
3.3 Performance of DoE design From previous experiments was known that the cultivation temperature L has an influence on the productivity because protease activity is inhibited at lower temperatures. Therefore an optimal temperature had to be found for lowest protease activity. The pH during cultivation could also change the protease activity and was under investigation as well. The influence of the inducer concentration cS2M (methanol) and of glycerol addition in an exponential feed profile with a set point µS1w had to be determined as two additional parameters. The search domain of optimal cultivation conditions is shown in tab. 1. The experiments were performed as sequential cycles in a repeated methanol fed batch with FM22 minimal media. Fig. 5: Automation structure of a DoE-research bioreactor A research bioreactor is build up with various measurement and control units as shown in Fig. 5. Besides the standard control task of maintaining a constant temperature, the pH is titrated with 12.5% ammonium hydroxide or 2 M phosphoric acid and the dissolved oxygen tension (pO2) is controlled to a set point of 25% in a cascade with agitation speed or with oxygen blending. The inducer concentration is measured inline with a diffusion probe (Kempe, Germany) and manipulated with a peristaltic methanol pump (Watson Marlow, UK). The turbidity (optek-Danulat, Germany) as well as the broth capacitance (Aber, UK) are measured inline to estimate the cell density and cell viability. An online off gas measurement of O2 and CO2 (BlueSens, Germany) provides the
After an automated harvest procedure the residual cells were diluted with fresh media and used as initial condition to grow on methanol in a new sequence. Each cycle started with a volume of 3.7 l and a cell density of 15 gl-1. The screening parameters were changed every sequence. Center point experiments were performed in the beginning, in the middle and at the end of the cultivation. A time course of these experiments is shown in Fig. 6. The total protein concentration cPtotM was determined with a Bradford Kit (BioRad, Germany). The target protein concentration cP1M was quantified from a SDS-PAGE using a fluorescent staining and a Multi-Imager (BioRad, Germany). During optimization of enzyme production a target protein activity measurement was not available.
Table 1: Search domain for parameter screening low
center point
high
L 18°C
22°C
26°C
pH 4.5
5
5.5
cS2M 0.5 gl-1
1.75 gl-1
3 gl-1
µS1 0 h-1
0.025 h-1
0.05 h-1 Fig. 6: Fully automated screening steps in a fractional factorial design
µS1w: set point of exponential glycerol feed, cP1M: target protein concentration (□), cPtotM: total protein concentration (◊), PRDj: target protein productivity
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3.4 Result of DoE screening The enzyme concentration cP1M and the cell density cXL were determined at the beginning (tj) and at the end (tj+1) of each experiment. The performance index of this cycle j is the mean productivity PRD, VM (t j1 ) cP1M (t j1 ) VM (t j ) c P1M (t j ) , PRD j (1) VL (t j1 ) t j1 t j which is related to the liquid volume VL at the end and the duration of this cycle. For a correct calculation of the mass of secreted protein, the volume of the media VM, c (t ) VM (t k ) 1 Z / X XL k VL (t k ) , k = j, j+1 Z
(2)
has to be calculated with the cell density cXL, using the weight ratio Z/X between wet and dry cells and the density Z of the wet cells. The experimental determined productivity with (1) revealed two significant and two non significant parameters with a response variability R2 of 0.8 and a predictive ability Q2 of 0.7. A significant influence on the expression of CALB was found in the inducer concentration cS2M (methanol) and in the cultivation temperature L. An exponential glycerol feed with its set point µS1w and a variation of the pH had no significant influence on the protein expression. In Fig. 7 the resulting response contour plot of the productivity is shown. Only the two significant parameters cS2M and L were plotted, whereas the expression rates are represented in colors from red (high productivity) to blue (low productivity). A high enzyme expression rate was found in the lower left corner (red) at a low temperature and a low methanol concentration.
A cultivation temperature below 18°C is not possible to maintain inside the DoE-research bioreactor. A methanol concentration of 0.5 gl-1 is the limit of quantification. Therefore this corner was taken as the optimal conditions for a high productivity. Using this knowledge the enzyme production was carried out as described below. 4. FULLY AUTOMATED PRODUCTION Each enzyme production cycle began with a transfer of 225 g cells from the cell breeding reactor. The broth was diluted in fresh media to a volume of 15 l resulting in a start cell density of 15 gl-1. The needed transfer volume was calculated with the cell density of the cell breeding reactor, estimated from the turbidity signal. A production cycle last 24 hours and was carried out at a cultivation temperature of 18°C, pH 5 and a controlled methanol concentration of 0.5 gl-1 for high expression rates as described above. No bi-substrate feed with glycerol was performed. In Fig. 8 nine production cycles are presented, where cycle 2.9 last 48 hours and shows no decrease of productivity in an extended expression time. Production cycle 2.5 follows a bad cell breeding cycle 1.5 with accidentally no pH control. This shows the dependency on a good cell breeding for high expression rates in the production. On the other hand the process strategy offers a complete regeneration of the product expression, as seen in the protein concentration chart in cycle 2.6. A new 12 hour batch phase with unlimited growth on glycerol in the cell breeding step probably causes this process recovery after a cycle with poor conditions.
Fig. 7: Response contour plot of parameter screening
Fig 8: Protein production cycles in the fully automated Integrated Bioprocess AP1M: UV absorption of enzyme concentration (∆), cS2M: inducer concentration (•), AP1actM: enzyme activity against tributyrin (□), pO2: dissolved oxygen tension
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4.2 Process automation with S88 The process automation was done with a S88 programming in the SCADA system MFCS/win 3.0 (Sartorius, Germany). A part of the block structure for the automation is shown in Fig. 9. These are the steps for an automated harvest of the production bioreactor into the disc clarifier, the refresh steps to fill fresh media into the bioreactor and the cell transfer steps from the cell breeding into the production bioreactor. This can be initialized via a manually changed variable value or by batch time.
100 mM sodium hydroxide (TitriPak, Merck, Darmstadt). One unit of lipase activity was defined as the formation of one µmol butyric acid per minute. 5.2 Atline target protein concentration The target protein concentration was measured every 35 minutes via HPLC (LaChrom Elite, VWR, Germany) at 20°C with a modified method from Trodler et al. A buffer exchange of the sample was performed with a size exclusion chromatography step using Sephadex G25 superfine in a 5 ml column (HiTrap Desalting, GE). CALB was bound (20 mM sodium citrate, pH 3) to a strong cation exchanger in a 1 ml column (HiTrap SP FF, GE) and eluted (30 mM sodium acetate, pH 5.5) with increasing the pH of the buffer. The absorption was measured at 280 nm.
Fig. 10: Atline measurements for lipase activity (up right) and concentration of the target protein (down right) 6. INTEGRATED PROCESSING Fig. 9: S88 programming structure for the initialization of a new production cycle
6.1 Cell harvest with disc clarifier
5. ATLINE PRODUCT OBSERVATION
The whole culture of the production bioreactor is clarified in the last stage via a semi-continuously disk clarifier SC1 with optimal parameters of 12500 rpm rotational speed, 40 lh-1 inlet flow and 3.4 bar outlet pressure.
As shown in Fig. 10, a cell free sample flow from the production bioreactor was carried out at 15 mlh-1 with a peristaltic pump through a filtration probe with a pore size of 0.2 µm (ESIP, Trace, Germany). The supernatant was collected in a vial, which level was maintained at 6 ml with an outlet flow. An automatic titrator dispersed the sample to the sample loop of a HPLC device and into the reaction chamber of the titrator after an optional dilution. 5.1 Atline activity measurement against tributyrin The hydrolytic activity of the enzyme against tributyrin was measured atline using an automatic titrator in pH-stat mode (ProcessLab 875, Metrohm, Germany). The reaction was carried out at 36°C and initiated by adding 100 µl of the sample to the substrate solution (5% tributyrin, 2% gum arabicum). The pH was controlled to 7.0 by titration with
For the optimization of the clarification the inlet flow and the outlet pressure were under research (data not shown). With a solid holding space of 0.7 l and an inlet flow of 40 lh-1 a cell discharge was done every six minutes followed by an approximately six minute recirculation until the cell separation was stabilized again. The clarification was satisfying because a reduction of solid content with a factor of 500 was reached, represented by the ratio between the turbidity of the inlet and outlet flow. In Fig. 11 one cycle of the whole Integrated Bioprocess is shown exemplarily. As presented in cycle 3.3 the enzyme activity AP1actC was measured in the clarifier outlet and in samples from the recirculation. No significant loss of product quality during the first down-stream step was determined.
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Fig. 11: One cycle of the whole Integrated Bioprocess
Ein, Eout: turbidity in clarifier inlet and outlet, AP1M: UV absorption of enzyme concentration (∆), AP1actK: activity of enzyme against tributyrin (□), Fout: outlet flow of the clarifier
7. OUTLOOK A further reduction of impurities from the supernatant such as cell debris or DNA could be done with a micro filtration. SDS-PAGE analysis showed already a high purity of the supernatant after clarification. For an industrial application of CALB no further purification is necessary. The next steps will be an ultra filtration to concentrate the supernatant and a lyophilization. With the opportunity to monitor the product quantity as well as the product activity in the production reactor and an additional application of these measurements in the down stream steps a fully automated plant offers a high potential to develop a self optimizing integrated production. This will be done with online applications of heuristic parameter search algorithms [Nelder and Mead, 1965] or genetic algorithm for parameter identification [Takors et al., 2001] in the future.
Eriksson, L., Johansson, E., Kettaneh-Wold, N., Wikström, C., Wold, S. (2008): Design of Experiments – Principles and Applications, MKS Umetrics AB, Umea, Sweden. Nelder, J.A., Mead, R. (1965): A simplex method for function minimization. Computer Journal (7), 308-313. Rotticci, D., Haeffner, C., Orrenius, C., Norin, T., Hult, K. (1998): Molecular recognition of sec-alcohol enantiomers by Candida antarctica lipase B. J. Mol. Catal. B. 5, 267-272. Takors, R., Weuster-Botz, D., Wiechert W., Wandrey C. (2001): A Model Discrimination Approach for Data Analysis and Experimental Design. In Hofman, M. and Thonart, P. (ed): Engineering and Manufacturing for Biotechnology, Vol. IV, 111-128. Kluver Academic.
REFERENCES
Trodler, P., Nieveler, J., Rusnak, M., Schmid, R.D., Pleiss, J. (2008): Rational design of a new one-step purification strategy for C. antarctica lipase B by ion-exchange chromatography. J Chromatogr A 1179, 161-167.
Anderson, E.M., Larsson, K.M., Kirk, O. (1998): One biocatalysis - many applications: The use of Candida antarctica B-lipase in organic synthesis. Biocatal Biotransform 16 (3), 181-204.
Uppenberg, J., Hansen, M.T., Patkar, S., Jones, T.A. (1994): The sequence, crystal structure determination and refinement of two crystal forms of lipase B from Candida antarctica. Structure 2 (4), 293-308.
Cregg, J.M., Vedvick, T.S., Raschke, W.C. (1993): Recent advances in the expression of foreign genes in Pichia pastoris. Bio/Technology 11, 905-910.
ACKNOWLEDGEMENTS
Degn, P., Zimmermann W. (2001): Optimization of carbohydrate fatty acid ester synthesis in organic media by a lipase from Candida antarctica. Biotechnol. Bioeng. 74, 483-491.
The authors thank the BMBF - German Federal Ministry of Education and Research for the financial support, Project BIOKATALYSE2021 - P13: FKZ 0315167, the Institute of Technical Biochemistry at the University of Stuttgart for providing the X-33 Pichia pastoris strain and industrial partners for their support.
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