Classifications

• • C—CHEMISTRY; METALLURGY
  • C12—BIOCHEMISTRY; BEER; SPIRITS; WINE; VINEGAR; MICROBIOLOGY; ENZYMOLOGY; MUTATION OR GENETIC ENGINEERING
  • C12P—FERMENTATION OR ENZYME-USING PROCESSES TO SYNTHESISE A DESIRED CHEMICAL COMPOUND OR COMPOSITION OR TO SEPARATE OPTICAL ISOMERS FROM A RACEMIC MIXTURE
  • C12P13/00—Preparation of nitrogen-containing organic compounds
  • C12P13/04—Alpha- or beta- amino acids
  • C12P13/22—Tryptophan; Tyrosine; Phenylalanine; 3,4-Dihydroxyphenylalanine
  • C12P13/222—Phenylalanine

Definitions

• the present invention relates to a process for in situ removal of a product in a bioconversion process.
• a bioconversion process may for instance be a fermentation process or an enzymatic process.
• Products formed in a bioconversion process may be toxic to the microorganism and/or may be inhibiting to a biocatalyst (enzyme or microorganism) when the product reaches a certain concentration.
• the product may be degraded or side reactions may occur in a bioconversion process. Therefore, it is difficult to reach high product concentrations in a bioconversion process.
• WO03/092853 discloses a method for the integrated separation of one or more organic substances present in an aqueous bio- process mixture containing at least one positively charged and/or chargeable group containing nitrogen by means of reactive extraction in at least one step.
• the reactive extraction applied in the process according to WO 03/092853 uses at least one liquid- liquid centrifuge. The isolation of products is often difficult.
• micoorganisms are grown in a medium consisting of numerous compounds. Products produced by the microorganisms need to be removed from the fermentation broth, which consists of liquid and solid compounds.
• the aim of the present invention is the provision of an alternative process for the in situ removal of a product in a bioconversion process.
• the present invention relates to a process for the in situ crystallisation of a product in a bioconversion process.
• the process comprises (i) preparing the product under bioconversion conditions, wherein the bioreactor comprises a reaction mixture which comprises (a) a solvent and the product; and (b) a catalyst for bioconversion, (ii) recycling part of the reaction mixture through an external crystallisation loop and (iii) recovering the product in crystal form by removal of the solvent.
• in-situ crystallisation is meant to include a process in which the product is removed during production and recovered in crystal form.
• in-situ crystallisation takes place in an external crystallisation loop, i.e. outside the bioreactor, and involves the removal of the solvent.
• the removal of solvent takes place before or during crystallisation.
• Any crystallisation method may be used for instance cooled crystallisation or uncooled crystallisation, such as evaporative crystallisation or anti-solvent crystallisation, in which process an anti-solvent is added which decreases the solubility of the product.
• the crystallisation may also take place at a temperature which is higher than the temperature of the bioconversion reaction, for instance when the product is L- phenylalanine.
• crystallisation takes place at a temperature of at least 30 degrees C, more preferably at a temperature of at least 35 or 40 degrees C.
• the product in crystal form is advantageously separated from the mother liquor by known methods in the art, for instance filtration or centrifugation.
• a first advantage of the process according to the present invention is that a higher production rate may be obtained in bioconversion processes due to the external crystallisation loop. Another advantage is that the solvent removed may be used to wash the crystallised product which reduces the amount of water to be used. Yet another advantage of the present invention is that the removal of solvent in the external crystallisation loop results in recycle flows with reduced volumes. As a consequence the recycle flows become more manageable.
• Another advantage of the process of the present invention is that the product is obtained in crystallised form which is the preferred form for most commercial chemicals which are produced by fermentation. In the process, the number of downstream processing steps are reduced and the production process of the product is thus simplified.
• Another advantage is that less equipment is required to obtain the product in pure crystalline form, which leads to a cost reduction.
• Yet another advantage is that the use of extraction solvents that are normally required to obtain the product in pure form, is avoided.
• a bioconversion process refers to a process in which one or more products are produced by means of a biological reaction.
• the catalyst used is typically an enzyme or a microorganism and may be free or immobilized.
• under bioconversion conditions refers to conditions which allow for bioconversion.
• the bioconversion process is a fermentation process and the catalyst is a microbial cell.
• micro- organisms that are used in industrial fermentation processes and that can be used in the process of the invention include Saccharomyces cerevisiae, Escherichia coli, Penicillium sp., Corynebacterium sp., Bacillus subtilis, Pseudomonas putida, Pichia pastoris, Streptomyces clavuligerus, and microalgae such as Dunaliella sp. Other micro-organisms can be used as well.
• Products which can be produced in crystal form by industrial fermentation processes and which are suitable for the present invention include amino acids, peptides, proteins, carboxylic acids, carboxylates, carbohydrates, nucleic acids, steroids, isoprenoids, carotenoids, cyclitols or any other primary or secondary metabolites, as well as derivatives, salts or esters thereof.
• Some specific examples include phenylalanine, fumaric acid, sodium glutamate, lysine, tyrosine, tryptophan, aspartate, penicillin, cephalosporin, riboflavin, clavulanic acid and ⁇ -carotene. This list of products merely serves as example and should not be considered limiting.
• ⁇ cells are grown at 15 to 40 0 C, preferably at about 30°C-37°C in appropriate media.
• Preferred growth media for the present invention are common commercially prepared media.
• Other defined or synthetic growth media may also be used and the appropriate medium for growth of the particular micro- organism will be known by someone skilled in the art of microbiology or fermentation science.
• the pH at the start of the fermentation is typically between pH 2.0 and pH 9.0, preferably between pH 3.0 and pH 8.0, more preferably between pH 5.0 and pH 8.0.
• Reactions may be performed under aerobic or anaerobic conditions. Preferably, the reactions are performed under aerobic conditions.
• Reactions that can be used in the method according to the present invention include for example the production of L- phenylalanine from sugars by recombinant Escherichia coli, the production of clavulanic acid by the filamentous microorganism Streptomyces clavuligerus, the production of glutamic acid by Corynebacterium glutamicum, and the production of carotenes by microalgae such as Dunaliella sp.
• the solvent removed will typically be aqueous since the fermentation medium will typically be aqueous, viz. water or a mixture of water and an organic solvent such as an alkane, alcohol, ether, or ketone, in a preferred embodiment, the solvent is water.
• the solvent may be removed by evaporation, dialysis, pervaporation, reverse osmosis, electrodialysis, nanofiltration, eutectic freezing or other known methods.
• the solvent removed in the crystallisation loop may be partly recycled to the bioreactor (see Figure 1 ) or may be used for washing the crystals.
• the present invention may be practised using batch, fed-batch or continuous fermentation.
• Fed-batch fermentations may for example be performed with carbon feed, for example glucose limited or in excess.
• the fermentation feed e.g. glucose or ethanol
• the fermentation feed can also partly or completely be fed to the crystalliser in the external crystallisation loop. Partly or completely feeding the fermentation feed to the crystalliser may reduce the solubility of the product (anti- solvent effect), so that a smaller proportion of the product will be in the mother liquor that is recycled to the bioreactor.
• Mother liquor, wash solvent and their mixtures can be purged to minimise the chance of product inhibition, or returned to the bioreactor to minimise losses.
• the process as described in the invention may be preceded by a growth phase in which the biomass is grown up to a suitable concentration, preferably 10-50 g cell dry weight per liter. Also, the process may be followed by a phase in which fermentation is halted, so that the remaining dissolved product can be crystallised and harvested.
• the external crystallisation loop preferably comprises means for separating the catalyst from the rest of the reaction mixture to obtain a catalyst-free solution.
• catalyst-free solution refers to a solution which is substantially free of catalyst.
• Suitable means for separating the catalyst include membranes, such as an ultrafiltration membrane, filters and centrifuges. Examples of catalyst-free solutions thus include permeates, filtrates and supernatants.
• solvent removal may be combined with anti-solvent addition, temperature reduction, pressure changes, pH changes or seed crystal addition to further improve the results and the process.
• the amount of removed solvent depends on the solubility of the product and on the specific needs. Preferably, an amount of solvent is removed that just brings the product in supersaturation. In this way, the crystallisation is carried out in a modest way so that the quality of the product increases.
• the solvent and catalyst-free solution of the external crystallisation loop preferably flow back to the bioreactor for volume control in the bioreactor and for preventing loss of unused feed components.
• Some desired crystalline products would give at fermentation conditions a crystal morphology which is undesired.
• the present invention enables to select different temperatures for the fermentation and crystallisation process. In this way, the invention provides a control over the crystal form of the product. For example, at 37°C L- phenylalanine crystallises as fine needles, which turn the solution into a gel-like substance that cannot be filtered. However, at 50 0 C, flake-like crystals are obtained, which can be processed easily.
• the fermentation medium contained glucose (15 g-L “1 ), MgSO 4 JI-I 2 O (3.0 glL), CaCI 2 .2H 2 O (0.015 g/L), KH 2 PO 4 (3.0 g/L), NaCI (1.0 g/L), (NH 4 ) 2 SO 4 (5 g/L), FeSO 4 JH 2 O / NaCitrate (0.1 125 / 1.5 g/L), thiamine.HCI (0.075 g/L), trace elements solution (1.5 ml/L) and antifoam (1 mL/L) in demineralized water, adjusted to pH 6.5 with 6% NH 4 OH.
• the trace element solution contained CaCI 2 .2H 2 O (10 g/L), FeSO 4 JH 2 O (10 g/L), AICI 3 .6H 2 O (2.5 g/L), MnSO 4 -H 2 O (2.5 g/L), CoCI 2 .6H 2 O (1.75 g/L), Na 2 MoO 4 .2H 2 O (0.5 g/L), ZnSO 4 JH 2 O (0.5 g/L), CuCI 2 .2H 2 O (0.25 g/L) and H 3 BO 3 (0.125 g/L) in 5M HCI solution. Fermentation with cell retention was tested with wild-type strain E. coli K12.
• LB medium For pre-cultivation, 100 mL of LB medium was inoculated with approximately 1 ml. stock culture, and incubated for approximately 14 h at 37°C and 220 rpm in a shake flask incubator. Fermentation was started by inoculating 1 L fermentation medium with 100 mL pre-cultivation medium in a 2 L jacketed fermenter with three baffles and two Rushton turbine impellers of six blades (Applikon, the Netherlands). The fermenter was sparged using pressurized air at a flow rate of 2.2 L-min "1 , controlled by a mass flow controller (Brooks, United States). The stirring rate was kept constant at 900 rpm.
• the pH was maintained at 6.5 by controlled addition of acid (4M H 2 SO 4 ) and base (6% v/v NH 4 OH).
• the temperature was maintained at 37°C.
• Silicone based antifoam (10% w/w) addition was controlled by a level sensor.
• the ultrafiltration membrane system which enables cell retention, consisted of a Masterflex peristaltic pump for recirculation (Cole Parmer, United States), an UFP-500- C-3MA hollow fiber membrane cartridge (GE Healthcare, Belgium) and a Masterflex peristaltic pump for the permeate (Cole Parmer, United States).
• the membrane effective area was 0.014 m 2 and the nominal weight cut off was 500 kDa.
• Pressure was build up in the system by clamping the retentate tube. Pressure gauges were placed at the membrane feed, retentate and permeate. For this experiment both retentate and permeate were returned to the fermenter.
• the fed-batch was initialized after about 6 hours by starting the glucose feed
• the feed rate was fixed at around 8.2 g glucose feed per hour until the cell concentration reached about 25 g/L (approximately 16 hours). At that point, the glucose feed rate was reduced to about 4.3 g glucose feed per hour.
• Cell retention was started with a permeate flow rate of 1 1 mL-min "1 after the glucose feed had been reduced. The fermentation with cell retention was run for about 24 hours. The cell mass and carbon dioxide yields on glucose were comparable to those of a fermentation without cell retention suggesting that the cell retention system did not affect the fermentation performance.
• Reagent grade L-phenylalanine (Sigma Aldrich) was used. The solutions with the concentrations described below were prepared by dissolving L-phenylalanine in demineralized water and adjusting the pH to 6.5 by addition of 6% NH 4 OH.
• This vessel supplied a 2 L stirred vessel (referred to as buffer vessel) which originally contained about 790 g of a L-phenylalanine solution of 12.5 g/kg. In both vessels the stirrer speed was set to 400 rpm and the temperature was controlled at 50 0 C.
• the reverse osmosis system consisted of a Sepa CF Il membrane module (GE Osmonics), a 3CP1231 stainless steel plunger pump (Cat pumps) and a Sepa CF Thin Film SE reverse osmosis membrane (GE Osmonics).
• the effective membrane area was 0.014 m 2 .
• the retentate was recycled to the buffer vessel and the permeate was collected in a separate vessel. Pressure was build up in the system by closing the retentate valve. The pressure was applied in three stages (800, 1000 and 1400 kPa).
• the L-phenylalanine concentration in the feed vessel increased linearly; after about 5.8 hours the mass in the feed vessel was 490 g with a concentration of 30 g/kg. At that point, the L-phenylalanine retention by the membrane was 96%. This shows that it continuous concentration of L-phenylalanine was possible.
• a 2 L stirred vessel containing a L-phenylalanine solution of 42.5 g/kg with an outflow rate of 5 g-min "1 was used to mimic the steady-state concentration in the buffer vessel from the reverse osmosis set-up described above.
• This vessel supplied a 1.5 L stirred vessel (referred to as crystallizer) originally containing 500 g of a L- phenylalanine solution of 44.3 g/kg and with an outflow rate of 5 g/min.
• the temperature was controlled at 50 and 45°C in the buffer vessel and the crystallizer respectively.
• the crystallization was started by adding to the crystalliser 0.52 g of seeds in slurry form.
• the seeds had been prepared by sieving reagent grade L-phenylalanine and collecting the fraction between 125 and 212 ⁇ m.
• the crystals were kept in the crystallizer by a filter installed in the outlet port.
• An average steady state dissolved L- phenylalanine concentration of 40.4 g/kg (supersaturation 1.04) was reached after 4 hours. At this concentration, the crystal production rate in the crystallizer was 0.84 g/h.
• a phenylalanine-producing strain for example, the genetically engineered E. coli W31 10-4pF20
• This strain is grown following the fed-batch protocol as described by (Takors, Biotechnol. Progr. 20 (2004) 57-64) with the following modifications: starting fermentation mass 1.5 kg; constant tyrosine feed at about 0.2 g/h until the cell concentration reaches about 10 g/kg; at this point, the tyrosine feed is reduced to about 0.04 g/h and a cell retention system as described in Example 1 , but with a permeate rate of about 2.3 g/min is started.
• the external loop consists of a concentration step (by reverse osmosis) as described in Example 2 and a continuous crystallizer as described in Example 3.
• the concentration step is required to achieve the degree of supersaturation necessary ( ⁇ 1 -2) for crystallization in a subsequent step.
• a higher degree of supersaturation is undesirable since this leads to spontaneous crystallization of phenylalanine monohydrate.
• a supersaturation of 1.04 is chosen.
• the permeate stream from the cell retention module is connected to the buffer vessel.
• the temperature in this vessel is maintained at 50 0 C.
• the concentration is started as described in Example 2, but in two pressure steps: 1200 kPa for a L-phenylalanine concentration in the buffer vessel lower than 30 g/kg and 2000 kPa for higher concentrations.
• the crystallizer When the concentration of phenylalanine in the buffer vessel reaches the required supersaturation (1.04), the crystallizer is filled with about 0.25 kg of the contents of the buffer vessel at a rate of about 1 g/min. At this point the crystallization as described in Example 3, but with an inflow/outflow rate of about 1 g/min is started.
• the crystals remain in the crystallizer and the mother liquor is recycled to the fermenter.
• L-phenylalanine crystals are produced at a rate of 0.7 g/h. This rate equals the L-phenylalanine production rate by the microorganism in the fermenter. This production rate results in a productivity of approximately 1.14 g/L/h which is much higher than the state of the art production rate of 0.67 g/L/h (Takors, Biotechnol. Progr. 20 (2004) 57-64) reached in a system without in situ crystallization by solvent removal.
• Example 2 A fermentation with cell retention is carried out with a genetically engineered L-
• E. coli 4pF26 which is a derivative of E. coli K12 which contains a chromosomal deletion delta (pheA tyrA aroF) regarding pheA (coding for chorismate mutase/prephenate dehydratase), tyrA (coding for chorismate mutase/prephenate dehydrogenase), and aroF (coding for the tyrosine-sensitive DAHP synthase (2-deoxy- D-arabino-heptusonate 7-phosphate).
• pheA tyrA aroF a chromosomal deletion delta
• the external loop consists of a concentration step (by nanofiltration) required to achieve the degree of supersaturation necessary ( ⁇ 1.2) for crystallization in a subsequent step.
• a higher degree of supersaturation is undesirable since it might lead to spontaneous crystallization of Phe monohydrate.
• a supersaturation of 1.04 is chosen.
• a buffering vessel (feed vessel) is implemented for ease of temperature and concentration control. This vessel supplies both the nanofiltration unit and the crystallizer.
• a permeate stream from the biomass retention module is connected to the feed vessel.
• This vessel is filled with a liquid volume at least equal to the total hold-up volume of the external loop and the liquid volume in the crystallizer (see Table 2).
• the temperature in this vessel is maintained at 50 0 C by means of a jacket connected to a heating/cooling water bath. Below this temperature phenylalanine might crystallize in the concentration loop, which is undesirable. On the other hand, recirculation at the pressure required for concentration results in a continuous increase in temperature, which might be detrimental for the membrane.
• the concentration step consists of a SEPA CFII module (GE Osmonics), a nanofiltration membrane (Thin Film HL, GE Osmonics) and a three-plunger pump (CAT Pumps) to provide the appropriate feed flow rate at 5 bar.
• the minimum flow rate (stream 3) should be 2 L/min in order to achieve a suitable cross flow velocity in the unit (1 m/s).
• the retentate (stream 5) is returned to the feed vessel.
• the permeate (stream 4) which still contains nutrients like glucose and tyrosine, is recycled to the fermenter.
• the crystallizer When the concentration of phenylalanine in the feed vessel reaches the required supersaturation (1.04), the crystallizer is filled up and 0.039 g seed crystals are added. Seed crystals are obtained by sieving phenylalanine anhydrate (Sigma Aldrich, purity >99.0%) and collecting the fraction between 90-212 micrometer. Before addition, a slurry is prepared by adding 1 ml. of cold water. The crystals remain in the crystallizer while the mother liquor is recycled to the fermenter.

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