Conf. CHISA'98, Praha 1998, Czech Republic, Lecture no Change of impellers design may improve an efficiency of yeast growth process

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1 1 Change of impellers design may improve an efficiency of yeast growth process A. Prell, K. Šolar 1, H. Šafář, I. Fořt 1, M. Sobotka Institute of Microbiology CAS, Vídeňská 1083, Praha 4, 1 Czech Technical University, Faculty of Mechanical Engineering, Technická 4, Praha 6 INTRODUCTION Important number of aerobic fermentation processes performed under test, pilot-plant and industrial conditions as a part of biotechnological productions is oxygen limited. The low oxygen solubility in aqueous phase under physiological temperatures and pressures (27-38 C, kpa) influences the rates of energetic and consequently biosynthetic metabolism of cultivated microorganisms significantly, and negatively. Standard stirred bioreactors, equipped by a trio of standard Rushton turbine impellers on the same shaft and by a lower aeration sparger ring offer some possibilities of design improvements to increase the aeration efficiency, especially as amount of oxygen not only supplied, but as an absorbed portion by cultivated cells, too. Using this way the higher growth rate of the microorganisms can be achieved. These improvements also may decrease an energetic input of the whole fermentation run. As a model process a cultivation of the yeast Torulopsis on ethanol as a sole source of energy and carbon has been chosen. This process is used for the production of yeast biomass containing a certain part of heavy metals, especially selenium and chrome. The dried metal enriched biomass serves as a main part of the dietetic preparative "DIASTABIL" (Štros F. et al., 1987) for the glucose level control in human blood. Ethanol as a substrate is used for the product purity, but the energetic metabolism of ethanol is advantageous for the aim of this work because of its strict dependence on the oxygen supplied. This relationship can emphasise differences in the efficiency of the fermentation apparatus. As a model reactor a standard stirred tank with the total volume of 75 litres (working volume 50 L) was employed. On the base of model experiments process parameters for a pilot-plant bioreactor (total 300 L, working volume 200 L) were estimated and some verifying experiments have been carried out. Further aim of this work has been to check up the possibility of a substitution of the pure ethanol by a mixture of waste before- and after runs (B.A.-runs) coming from a distillation of fermentative alcohol. The experiments with B.A.-runs were carried out in parallel. EXPERIMENTAL The model bioreactor was CSTR type, 75 L of total volume (Bioengineering, Switzerland), made of stainless steel. It is furnished by a bottom lower powered shaft with a

2 2 mechanical seal and three high-speed impellers. The aeration sparger ring is located under the lower impeller and the mixing system is equipped with four radial wall baffles (Fig. 1). The pilot-plant bioreactor (300 L, made by the same manufacturer) is designed similarly, the dimensions are standardised. Significant dimensions of both vessels are mentioned in Tab. I. Fig. 1: Model bioreactor. Location of the three impellers on the same shaft avoids to a mutual influence of the impellers in both sizes of bioreactors (Fořt I. et al, 1989). For some experiments a configuration of the impellers was changed: the upper two Rushton turbines were substituted by two pitched blade impellers with six plane blades pumping downwards. The fermentors are driven by analogue control units that manage monitoring and control of basic process parameters as ph, temperature, impeller frequency of revolution (RPM, air flow, DOT, output pressure. Moreover, these local units are connected with a central monitoring and control system that ensures archiving the process data and higher types of Tab. I: Geometric dimensions of bioreactors. Total height Inner diameter Impeller rotor diameter Height of the first impeller above the bottom Height of the second impeller above the bottom Height of the third impeller above the bottom Total volume Initial volume of the batch Final volume of the batch Width of the baffles Aeration sparger diameter Aeration pipe diameter Height of the level (final) sym H F D d H 1 H 2 H 3 V T V p V k b D v D t H unit L L L 75 L L

3 3 control (Prell A., 1996). This system also includes a measurement of the oxygen and carbon dioxide contents in a gas-off (Hartmann-Braun infrared and paramagnetic analysers, Germany). The microorganism used, the yeast Torulopsis ethanolitolerans, is adapted on higher concentrations of ethanol in a medium (up to 15 g/l). Mineral medium: H 3 PO 4 (75%) 2.46 kg, KOH 1.8 kg, MgSO 4 x7h 2 O 1.92 kg, ZnSO 4 x7h 2 O 30 g, CuSO 4 x5h 2 O 1.2 g, MnCl 2 x4h 2 O 3 g, FeSO 4 x7h 2 O 36 g, (NH 4 ) 2 SO g per 1000 L of medium. Inoculum was prepared in a laboratory fermentor from an agar culture of the strain on ethanol under the constant temperature and ph (30 C, 3.8). Initial concentration of biomass in the model tank (50 L) was 3 g/l, in the pilot-plant one (200 L) 4 g/l. Fermentation runs in both tanks were also led under constant conditions, but the impeller frequency of revolution and air flow rate were increased gradually up to the maximum values for the given experiment depending on the growth of culture and its oxygen demands. Ethanol as a toxic substrate must be fed stepwise during fermentation - the feeding was realised as a task of the control system according to the actual need of the culture (Prell A. et al., 1997). Biomass dry weight was determined by drying up to a constant weight. The total dose of ethanol (B.A.-runs) was 5 L for the model tank and 20 L for the pilot-plant one. Apparent concentration of ethanol in B.A.-runs was determined by means of an alcoholometer. All fermentation runs under the same conditions were performed at least two times. RESULTS AND DISCUSSION Evaluation of the fermentation runs in the model and pilot-plant bioreactors was based on several global indicators. At first, for the growth efficiency the X/S yield was used: Y x/s X S, (1) where delta X is the total amount of biomass producted and S is the total amount of substrate consumed (calculated on absolute ethanol). Next indicator was the productivity of the system, Y x/s related to the total time of cultivation: P Y x/s t X S. t. (2) Balance between the cell catabolism and anabolism can be under these conditions (ethanol as a sole energy and carbon source) characterised by respiratory quotient: RQ dco 2/dt do 2 /dt, (3) i.e. the ratio of carbon dioxide evaluated to oxygen consumed. This indicator has been evaluated in the second phase of fermentations, only, when the maximum aeration conditions has been reached and flows of both gases in the gas-off occurred in dynamic steady-state. The higher values of RQ point to a lower utilisation of carbon for the biomass synthesis, the energetic metabolism is less efficient. The yield of oxygen consumed to oxygen supplied was expressed by the ratio:

4 4 Y O2 n O2in n O2out n O2in, (4) where n O2x (x=in,out) are the total molar amounts of oxygen produced/supplied during the whole fermentation run. For example: n O2in x t O2in.Q t gin.dt, (5) where Q gin is the inlet gas flow rate, x O2in is the molar (volume) concentration of oxygen in the inlet gas. Economic efficiency of the processes was expressed by means of mixing expenses N M, energetic expenses N E and total expenses N T. The energetic expenses on the whole run included the expenses on mixing (power input), aeration (power input of compressors) and cooling. The total expenses involved the prices of substrates as well. All the expenses are relative to the weight unit of biomass produced. Influence of the impeller change in the model bioreactor A typical course of the fermentation run in a model reactor is depicted on Fig. 2. Fig. 2a: DOT, RPM, air flow rate, model tank. DOT AirFlowRate [l/min] (in Thousands) DOT RPM=850 1/min RPM 0.6 AirFlow Rate time [h] Fig. 2b: O2, CO2 in gas-off, model tank. O2 [%] CO2 [%], RQ O2 18 CO RQ time [h]

5 5 Fig. 2c: Ethanol feeding, model tank. FeedRate [L/h] time [h] The process can be divided into two phases: During the first phase the oxygen supply was sufficient, the concentration of biomass did not reach the limit value and aeration and mixing have increased gradually. When desired air flow rate and frequency of impeller revolution (80 L/min and 850 min -1, in the working volume 50 L) were achieved the culture got under the oxygen limit and the actual dissolved oxygen tension (DOT) has drawn near zero. This is the so called second phase of fermentation when the ratio RQ is steady. Tab. II: Growth efficiency indicators. reactor impeller substr. Y x/s std. P std. RQ std. YO2 std. Y type type type dev. [g/(l.h)] dev. dev. dev. Co/Ci 75 L radial ethanol L axial ethanol L radial ethanol L radial B.A.-runs L axial B.A.-runs L radial B.A.-runs The results obtained when the upper two standard Rushton turbines (radial impellers) were changed by two pitched blades impellers pumping downwards (axial impellers) are shown in the Tab. II. and III. Tab. II shows that the yields Y x/s of the whole runs are still the same when consider the standard deviation of both results. So the impeller type does not influence the total yield. However, the lower mixing intensity extends the fermentation time so that the productivity results to be lower (-14%). The values of RQ seem to be close, the influence of impellers on energetic coupling is not significant. The oxygen yield Y O2 reached higher value for radial impellers. Tab. III displays an overview of the economic expenses of the particular types of fermentation runs. Mixing expenses N M in case when axial impellers have been used lower significantly (-26%), but the energetic ones N E slightly rise - the influence of a longer cooling. Effect of substrate costs reduces this difference in the total expenses N T. The costs are relative

6 6 to the weight unit of biomass produced. Tab. III: Economic efficiency indicators. reactor impeller substr. NM std. NEstd. NT std. type type type CZK/kg dev. CZK/kg dev. CZK/kg dev. 75 L radial ethanol L axial ethanol L radial ethanol L radial B.A.-runs L axial B.A.-runs L radial B.A.-runs Estimation of process parameters for the pilot-plant bioreactor When the model tests had been performed we calculated the estimation of aeration conditions for the pilot-plant reactor (working volume 200 L). Probably the most suitable scale-up criterion appears to keep constant VVM. Unfortunately, the air flow rate 320 L/min which is needed for the VVM = 1.6 as in the case of the model tank, was not possible to reach for technical limitations of the pilot-plant bioreactor. Another model condition was chosen: the constant superficial gas velocity w 0. w 0 4Q gin D 2, (6) where Q gin is input gas (air) flow rate and D is the inner diameter of the bioreactor. The average value of w 0 for the model tank was m/s thus air flow rate for the 300 L tank was determined 168 L/min. As a condition for the equivalent intensity of mixing is considered the constant specific aerated impeller power input P g /V. Quantity P g was estimated from the dimensionless power number P o = for three radial impellers and P o = 7.94 for combination of two axial and one radial rotors (Bujalski W. et al., 1993): P o P, n 3 d 5 (7) where P is ungassed power input of all impellers on the same shaft, ρ is density of the liquid (~1000 kg/m 3 ), n is the impeller frequency of revolution and d is impeller diameter. Value of the ungassed power input was used for calculation of the gassed one (P g ) according to Hughmark correlation (Hughmark G.A., 1980): P g P C. V g nv n 2 d 4 ghv 2/3 0.2, (8) where V is volume of liquid phase in a bioreactor, g is gravitational acceleration, and h is the impeller height (the other parameters were mentioned in previous relations). The constant C was used from Fořt measurements (Fořt I. et al., 1993, model tank C = 0.13, pilot-plant tank C = 0.11). The ratio P g /V for the model tank (radial impellers) was calculated to be equal kw/m 3 and the equivalent impeller frequency for the pilot-plant tank was estimated

7 7 n = 580 min -1. The courses of P g /V are depicted on Fig. 3. Fig. 3: Time courses of Pg/V. Pg/Vw [W/m3] (in Thousands) time [h] rad50 ax50 rad200 Pilot-plant bioreactor tests For verification of validity of the scale-up relations it was not possible to keep the constant VVM in the pilot-plant reactor. Although the model condition of constant superficial gas velocity and specific power input had been applied (Fig. 3), a lowered aeration capacity resulted in the lower productivity of pilot-plant bioreactor (see Tab. II). Extending the fermentation time and a higher level of the oxygen limit led to lowering the energetic efficiency of the cell metabolism. It might be demonstrated on an increasing of RQ and principally on lowering the production yield. Moreover a longer fermentation time could enhance an ethanol loss by airing. It is interesting that the oxygen yield Y O2 shows higher values for the pilot-plant tank than these for the model one. The lower rate of the oxygen supply probably leads to its better utilisation though the average growth rate has decreased. A comparison of expenses expresses mixing expenses N M for the pilot-plant runs exceed the model ones by approx. 65%, while the energetic costs are by 13% (the cooling is by the pilot-plant reactor more effective). At total expenses N T this discrepancy is reduced by the dominant impact of the substrate cost (see Tab. III). Influence of the distillery before- and after- runs In the frame of technology-improvement experiments the substitution of ethanol substrate by the mixture of distillery before- and after-runs (B.A.-runs) was testified. For the model experiments B.A.-runs type 1 (88% of apparent alcohol) and B.A.-runs type 2 (68% of apparent alcohol) during the pilot-plant trials have been used - we had failed when managed a sufficient amount of the alternate substrate. B.A.-runs contain besides ethanol lighter and

8 8 heavier alcohols and relative compounds (type 1 (g/l): acetaldehyde 10.1, methanol 0.8, ethylacetate 11.4, n-propanol 9.2, iso-butanol 14.4, iso-amylalcohol 15.4). It might be expected that the lighter compounds are deaired iediately while the heavier alcohols are utilised when ethanol is exhausted. A content of other composites, according to a presumption, had decreased the energetic efficiency of cell growth - the yield and the productivity were reduced (-8.5% and 15%, see Tab. II). A higher consumption of cell energy is the reason of the strong heightened RQ (20%) while oxygen need appears to be similar (Y O2 ). Another situation came using B.A.-runs type 2: this substrate turned up to be slightly more suitable for cell utilisation because all the indicators result similarly or even better in the case of the alternate substrate. Considering the standard deviations, the differences are neglectable. Production energetic costs confirm previous "efficiency" results (see Tab. III): B.A.-runs type 2 appears to be more profitable in the pilot-plant bioreactor, B.A.-runs type 1 raises the mixing and energetic costs in the model tank. Total costs recoend to use the alternative substrate because of the price difference (the model -30%, the pilot-plant -27%). The lowering of productivity and yield do not suffice to cover the advantage of a lower price. Carbon balance of the fermentation process Ethanol oxidation and biomass building-up can be expressed by a suary balance equation: C 2 H 5 OH O 2 CH 1.72 N 0.15 O 0.41 CO 2 H 2 O (9) where approximate formula CH 1.72 N 0.15 O 0.41 expresses an elementary composition of a yeast biomass grown on ethanol (Battley E.H., 1960). According this equation molar amounts of carbon are equivalent. For the carbon balance calculation the suary reaction equation (9), data ethanol of added and data of CO 2 evaluated in gas-off have to be used. Density of pure ethanol was kg/m 3, concentration 99% vol., reaction input volumes were 5 L for the model tank and 20 L for the pilot-plant one. In case of B.A.-runs an apparent concentrations of ethanol were used. The amount of carbon dioxide evaluated was calculated from the measured content of this gas in the gas-off. Output gas flow rate have not been measured, then we must express this quantity from balance of the inert gas (nitrogen): x N2in.Q gin x N2out.Q gout, x CO2out x O2out x N2out 1, x CO2in x O2in 1, (10) (11) (12) Q gout 0.79.Q gin 1 x CO2out x O2out, (13)

9 9 n CO2out x t CO2out.Q t gout.dt, (14) where x N2in, x N2out, x O2in, x O2out, x CO2in, x CO2out are molar (volume) fractions of particular gas in the input, resp. output flow. Q gin, Q gout are the total flow rates of input and output gas. Constant corresponds to the number of litres in one mole of a ideal gas (30 C), value of 0.79 is the volumetric content of nitrogen in air. As seen from Tab. IV, relative deviations between the carbon output and input in case of 75 L fermentor do not exceed 15% and in the 300 L one are lower significantly, up to 5%. So it may be concluded that the relative error of measurements is reduced in the bigger tank, where the air flow rate is approx. two times and the batch volume is approx. four times higher. Tab. IV: Carbon balance, examples. reactor impeller substr. biom. C in O2 in C biom CO2 o diff. C rel. diff type type type [g/l] [mol] [mol] [mol] [mol] [mol] C [%] 75 L radial ethanol L axial ethanol L radial ethanol L radial B.A.-runs L axial B.A.-runs L radial B.A.-runs It is also interesting that all the deviations were positive (for all the experiments, not included in the Table). This finding may signify a permanent error of gas-off measurements, or, more probably, a considerable variance in a determination of the biomass elementary composition. Average values of Y Co/Ci confirm a trend of Y x/s values (see Tab. II). Oxygen balance in the second fermentation phase Our fermentation experiments were carried out in fed-batch regime, when ethanol was added continuously according to the growth need. So oxygen became the limiting element of the growth. During the second phase of a fermentation run, when the constant aerating conditions had been set and DOT had been negligible, we used the gas-off measurements to try to make the oxygen balance and to find a relation between equipment scale-up and the oxygen supply. For the balance in the considered fermentation phase we supposed: DOT (c l ) comes near zero, x O2in = 0.21, x N2in = 0.79 for the input air (i.e. CO 2 content in the input air is insignificant, temperature and pressure are constant, K l a is constant for both bioreactors when the superficial gas velocity and P g /V are kept the same, the amount of the biomass produced (m b ) is in direct proportion to the amount of oxygen absorbed and to the length of the fermentation. The amount of oxygen absorbed is equal to the difference between its input and output:

10 10 Q O2 Q N2 X O2in X O2out, (15) where Q N2 Q gin 1 x O2in, (16) and X O2in x O2in 1 x O2in, X O2out x O2out 1 x O2out x CO2out. (17,18) Quantities X yin,out are relative molar fractions of gases (y = N 2, O 2, CO 2 ), Q gin is the total input gas flow rate and Q j (j = N 2, O 2 ) are partial gas flow rates. Balance of the oxygen becomes then form: dc l dt V V.K la c g c l. (19) For c l = 0 we may put down: Q O2 V.K l a.c g (20) and in combination with the above written oxygen balance we have: Q N2 X O2in X O2out V.K l a.c g (21) Here quantity c g* is the oxygen concentration in a liquid phase in equilibrium with the concentration in a gaseous phase, c l is the actual concentration of dissolved oxygen in a liquid and V is the volume of the liquid phase. Total amount of the biomass produced (m b = V.c b ) is considered to be directly proportionate to the oxygen absorbed, so we have: t f.q N2 X O2in X O2out k.v.c b, (22) or after rearranging: t f Q N2 V X O2in X O2out k.c b. (23) When we define the quantity "VVM" as: Q gin V VVM (24) and consider equation (16) we have finally: t f.vvm X O2in X O2out k 1 x O2in c b K.c b, (25)

11 11 where K is a new defined constant according to the relation (25) when gas volumetric ratio of oxygen in air x O2in is constant (0.21) and t f is fermentation time. When we denote results from the model bioreactor by index 1 and those from the pilot-plant tank by number 2, we can express the ratio fitting the equation (25): c b1 c b2 t f1.vvm 1 X O2in X O2out 1 t f2.vvm 2 X O2in X O2out 2. (26) Testing of an equality of this ratio proves the reliability of the constant K for both bioreactors, the linear dependency of the biomass growth on the oxygen supply and correctness of the VVM values for scale-up of the process considered. The results of the testing made under the above mentioned experimental (model and pilot-plant) conditions are suarised in Tab. V. Tab. V: Oxygen balance, eq. 26. substr. Left Right diff. diff. type side side L-P [%] ethanol B.A.-runs Differences 6 and -3.5% seem to be in very good agreement with the theory derived and confirm linear relation between the cell growth rate and oxygen supply rate. And, simultaneously, they approve validity of the constant K for both types of bioreactors hence the VVM can be employed to the process scale-up. CONCLUSIONS The first aim of this study was to investigate possibilities of design improvements of the standard stirred bioreactor for highly oxygen demanding fermentations by a substitution of impellers. The second objective was to transfer the successful changes to a pilot-plant bioreactor and verify the results on the same process. Finally, the third target proved the chance to use an alternative substrate the mixture of distillery before- and after- runs in both scales for the process studied. It may be concluded that: Efficiency of an employment of the carbon substrate (yield, Y x/s ) and the oxygen supplied had not been influenced when a combination of one radial and two axial impellers was used. At the same time, the mixing expenses were reduced because the power input of axial impellers is lower. On the other hand, the productivity was decreased and the energetic expenses rose trough a fermentation time lengthening. Scale-up brought not only a further reduction in the productivity, but the biomass yield was reduced, too. An employment of oxygen was better, but the supply was insufficient and inadequate to that in the model bioreactor the maintaining of P g /V and superficial gas velocity is not satisfactory rule for the scale-up of the process. To keep the VVM is probably necessary. The alternative substrate B.A.-runs can be used. Although the inner energetic efficiency of the cells had been worsened (RQ has risen), a difference in substrate costs covers all losses. Even the substrate type 2 exhibited the better results as ethanol.

12 12 Verification of the carbon and oxygen balances suggested confirms a reliability of all analytical methods used during the process of fermentation studied. The findings may form a base for further efficiency improvement. From the results presented several ways to optimise the process may be employed. When the axial impellers are used, the mixing energy will be saved, but the time will be lengthened. The alternative substrate saves the total expenses, but worsens the carbon employment. The smaller oxygen supply improves its efficiency, but decreases the growth rate. For further experiments an increasing of the frequency of impeller revolution would be examined to shorten the fermentation time (Arjunwadkar S.J. et al., 1998). REFERENCES Arjunwadkar S.J., Saravanan K., Pandit A.B., Kulkarni P.R.: Optimizing the Impeller Combination for Maximum Hold-Up with Minimum Power Consumption. Biochem. Eng. J. 1, 25-30, Battley E.H.: Growth-Reaction Equations for Saccharomyces cerevisiae. Physiologica Plantarum 13, , Bujalski W., Nienow A.W., Chatwin S., Cooke M.: Power Input of Rushton Disk Turbine Impeller. Chem. Eng. Sci. 42, 2345, Fořt I., Hájek J., Machoň V.: Energetic Efficiency of Two Impellers on the Same Shaft in a Cylindrical Baffled Vessel of High Height/Diameter Ratio. Collect. Czech. Chem. Coun. 54, , Fořt I., Hába Z., Sobotka M.: Power Input in Gas - liquid with Multiple Impellers. 11th International Congress of Chemical and Process Engineering CHISA'93, Prague, Czech Republic Hughmark G.A.: Power Input of Aerated Impeller. Ind. Eng. Chem. Proc. Des. Dev. 19, 638, Prell A.: A New Control System of Bioreactors in Pilot-Plant Scale - the Use of an EasyMAP Platform. 12th International Congress of Chemical and Process Engineering CHISA'96, Prague, Czech Republic Prell A., Šafář H., Sobotka M.: Optimisation of Yeast Biomass Production by Means of Control of Inflow of Toxic Source of carbon and energy (in Czech). 42th National Congress of Chemical and Process Engineering CHISA'97, Srní Štros F., Rut M., Adámek L., Beneš B., Táborský O.: Method of the yeast biomass preparation containing a high amount of active chrome complexes biologically bound (in Czech). Czech patent No (PV ), 1987.