Simulation and optimization of fed-batch culture for ethanol production from molasses

Two simulation methods for ethanol production from molasses by a flocculating yeast, Saccharomyces cerevisiae AM12, were investigated and molasses feeding was optimized. The first method was based on a deterministic model with fixed kinetic parameters and the second was based on regression analysis. The amount of ethanol produced in a fed-batch culture with multiple additions of molasses was simulated by both of these two methods. Simulated results of a fed-batch culture were compared with those of a simple batch culture by a model of regression analysis. The intermittent addition of molasses gave better production than a single addition at the beginning; more frequent addition may further improve production. The experimental results suggested the same. The effect of the amount of the added molasses on ethanol production was investigated by simulation. Repeated batch culture with and without intermittent addition of molasses in each batch was also done.List of Symbols C _e deviation of calculated results from experimental results - F m³ volume of feed medium added to the fermentor - P kg/m³ concentration of ethanol - P _M kg total amount of ethanol - S kg/m³ concentration of sugar - S ₀ kg/m³ concentration of sugar in the molasses feed medium - S _M kg total amount of sugar - V m³ culture volume - X kg/m³ concentration of cells - X _M kg total amount of cells - x _c calculated data - x _e experimental data - h^–1 specific rate of growth - kg-sugar/(kg-cell h) specific rate of sugar consumption - kg-ethanol/(kg-cell h) specific rate of ethanol production     

Integration of continuous production and recovery of solvents from whey permeate: use of immobilized cells of Clostridium acetobutylicum in a flutilized bed reactor coupled with gas stripping

An investigation was performed into the operation of an integrated system for continuous production and product recovery of solvents (acetone-butanol-ethanol) from the ABE fermentation process. Cells of Clostridium acetobutylicum were immobilized by adsorption onto bonechar, and used in a fluidized bed reactor for continuous solvent production from whey permeate. The reactor effluent was stripped of the solvents using nitrogen gas, and was recycled to the reactor. This relieved product inhibition and allowed further sugar utilization. At a dilution rate of 1.37 h^–1 a reactor productivity of 5.1 kg/(m³ · h) was achieved. The solvents in the stripping gas were condensed to give a solution of 53.7 kg/m³. This system has the advantages of relieving product inhibition, and providing a more concentrated solution for recovery by distillation. Residual sugar and non-volatile reaction intermediates are not removed by gas stripping and this contributes to high solvent yields.List of Symbols C kg/m³ Lactose concentration in reactor effluent - C _b kg/m³ Lactose concentration in bleed stream - C _c kg/m³ Lactose concentration in whey permeate feed - C _i kg/m³ Lactose concentration at reactor inlet - C _p kg/m³ Lactose concentration in condensed solvent stream (=0) - C _r kg/m³ Lactose concentration in recycle line (C _b=C _r) - C kg/h Amount of lactose utilized during certain time period - D h¹ Dilution rate of reactor, F _i/D=F/D - F dm³/h, m3/h F _i = Rate of feed flow to the reactor - F _b dm ³/h, m³/h Rate of bleed - F _c dm³/h, m³/h Rate of feed of whey permeate solution - F _p dm³/h, m³/h Rate of concentrated product removal - F _r dm³/h, m³/h Rate of recycle of stripped effluent to the reactor - P _l % Percent lactose utilization - R _l kg/(m³ · h) Overall lactose utilization rate - R _p kg/(m³ · h) Overall reactor (solvent) productivity - R _sl kg/h Rate of solvent loss - S kg/m³ Solvent concentration in reactor effluent - S _b kg/m³ Solvent concentration in bleed - S _c kg/m³ 0; Solvent concentration in concentrated whey permeate solution - S _i kg/m³ Solvent concentration at inlet of reactor - S _p kg/m³ Solvent concentration in concentrated product stream - S _r kg/m³ Solvent concentration in stripped effluent, S _r=S_b - S kg/h Amount of solvent produced from C amount of lactose in a particular time - ds/dt kg/(m³ · h) Rate of accumulation of solvents in the stripper - t h Time - V dm³, m³ Total reactor volume - V ¹ dm³, m³ Liquid volume in stripper - Y _P/S Solvent yield     

Penicillin acylase production by E. coli

Enzyme production with E. coli ATCC 11105, in a complex medium using phenylacetic acid as inducer is carried out in a stirred-tank reactor of 10 dm³ and an airlift tower-loop reactor of 60 dm³ with outer loop at a temperature of 27 °C. The optimum inducer concentration was 0.8 kg/m³, which was kept constant by fed-batch operation. The optimum of the relative dissolved O₂-concentration with regard to saturation is below 10% in a stirred-tank reactor and at 35% in a tower-loop reactor. It was kept constant by parameter-adaptive control of the aeration rate. In a stirred-tank enzyme productivity is slightly higher than in a tower-loop reactor, and much higher than in a bubble column reactor.List of Symbols CPR kg/(m³ h) CO₂-production rate - OTR kg/(m³ h) O₂-transfer rate - OUR kg/(m³ h) O₂-utilization rate - PAA phenylacetic acid (inducer) - RQ = CPR/OUR respiratory quotient - X kg/m³ cell mass concentration - _m h^–1 maximum specific growth rate     

Ethanol production with Zymomonas mobilis with synthetic medium in batch operation

Ethanol was produced with Zymomonas mobilis Z6 (ATCC 29191), in batch culture with synthetic medium on glucose as substrate and in the presence of aspartate. The concentrations of glucose, phosphate, ammonium, ethanol and dissolved O₂ and CO₂ in the medium and O₂ and CO₂ in the outlet gas as well as the cell mass by culture fluorescence were measured on-line. Cell mass, glucose and aspartate concentrations were measured off-line. In the presence of a sufficient amount of aspartate, the ethanol inhibition effect can be reduced considerably. However, the improvement with yeast extract is more incisive. The relationship between the intensity of culture fluorescence and cell mass concentration is linear, if sufficient aspartate is present.List of Symbols ASP kg/m³ aspartate concentration - CTR kg/(m³ · h) CO₂ transfer rate - N, NH₄ kg/m³ nitrogen concentration from NH ₄ ⁺ - P kg/m³ product (ethanol) concentration - p% product (ethanol) yield - PO₄ kg/m³ phosphate concentration - Q _E kg/(kg · h) specific ethanol production rate - kg/(kg · h) specific nitrogen uptake rate from NH ₄ ⁺ - Q _P kg/(kg · h) specific phosphate uptake rate - Q _s kg/(kg · h) specific substrate (glucose) uptake rate - S kg/m³ glucose concentration - S _O kg/m³ initial glucose concentration - Y _x/s kg/kg yield coefficient - h^–1 specific growth rate     

Production of glutamine by micro-gel bead-immobilized Corynebacterium glutamicum 9703-T cells in a stirred tank reactor

The effectiveness of using micro-gel bead-immobilized cells for aerobic processes was investigated. Glutamine production by Corynebacterium glutamicum, 9703-T, cells was used as an example. The cells were immobilized in Sr-alginate micro-gel beads 500 m in diameter and used for fermentation processes in a stirred tank reactor with a modified impeller at 400 min^–1. Continuous production of glutamine was carried out for more than 220 h in this reactor and no gel breakage was observed. As a result of the high oxygen transfer capacity of this system, the glutamine yield from glucose was more than three times higher, while the organic acid accumulation was more than 24 times lower than those obtained with 3.0 mm-gel bead-immobilized cells in an airlift fermentor under similar experimental conditions. During the continuous fermentations there was evolution and proliferation of non-glutamine producing strains which led to a gradual decrease in the productivity of the systems. Although a modified production medium which suppresses cell growth during the production phase was effective in maintaining the productivity, the stability of the whole system was shortened due to high cell deactivation rate in such a medium.List of Symbols C kg/m³ glutamine concentration - C _A mol/m ³ local oxygen concentration inside the gel beads - C _AS mol/m ³ oxygen concentration at the surface of the gel beads - De m²/h effective diffusion coefficient of oxygen in the gel bead - DO mol/m³ dissolved oxygen concentration - F dm³/h medium flow rate - K h^–1 glutamine decomposition rate constant - Km mol/m³ Michaelis Menten constant - QO _2max mol/(kg · h) maximum specific respiration rate - R m radius of the gel beads - r m radial distance - t h time - V _C dm ³ volume of the gel beads - V _L dm ³ liquid volume in the reactor - Vm mol/(m³ · h) maximum respiration rate - X kg/m³ cell concentration - x r/R - y C _A /C_AS - h^–1 cell deactivation rate constant - Thiele modulus defined by R(Vm/De Km) ^1/2 - C _AS /Km - _C kg/(m³-gel · h) specific glutamine formation rate - _c dm³-gel/dm³ V _C /V _L      

Mathematical modeling of production of microbial lipids

In the microbial lipid production system using the yeast Rhodotorula gracilis, CFR-1, kinetics of lipid accumulation and substrate utilisation at initial substrate concentrations in the range of 20–100 kg/m³ were investigated using shake flask experiments. A mathematical representation based on logistic model for biomass and Luedeking-Piret model for lipid accumulation gave reasonably good agreement between the theoretical and experimental values for substrate concentration less than 60 kg/m³. The kinetic expressions and parameters obtained through shake flask studies were directly applied to experiments in the laboratory fermentors also and the models were found to hold good for the prediction of the change of biomass, product as well as substrate with time. The attainment of a saturation in the intracellular lipid accumulation with time, however, was not predicted by the model which was shown to be an inherent feature of the Luedeking-Piret model.List of Symbols S ₀, P ₀ kg/m³ Initial concentrations of sugar and lipid respectively - S, S(t) kg/m³ Concentrations of sugar and lipid respeclively at any timet - p,p(t) L kg/m³ Maximum concentration of lipid produced - E % Maximum sugar utilised - dP/dt kg/(m³ · h) Rate of lipid production - -dS/dt kg/(m³ · h) Rate of sugar utilisation - _max h^–1 Maximum specific growth rate - X _max kg/m³ Maximum biomass reached in a run - P _max kg/m³ Maximum product concentration - m, n Constants used in Luedeking-Piret model in eq. (7) - , Constants used to predict residual sugar - k _e maintainance coefficient - Y _x g/g Biomass yield based on sugar consumed - Y _p g/g Lipid yield based on sugar consumed - (dP/d t)_stat kg/(m³ · h) Rate of lipid production at stationary phase - (dS/dt)_stat kg/(m³ · h) Rate of sugar utilisation at stationary phase     

Glutamic acid production in an airlift reactor with net draft tube

Cultivation of Brevibacterium divaricatum for glutamic acid production in an airlift reactor with net draft tube was developed. Cell concentration gave an index for adding penicillin G. On-line estimation of total sugar concentration yielded an identified model which was used for determination of the substrate addition. Fermentation for glutamic acid production requires high oxygen concentration in the broth. The proposed reactor has the capability to provide sufficient oxygen for the fermentation. Since the reactor is suitable for fed-batch culture, the cultivation of B. divaricatum for glutamic acid production in the proposed reactor is successfully carried out.List of Symbols a system parameter - b system parameter - C _c,in mole fraction carbon dioxide in the gas inlet - C _c,out mole fraction carbon dioxide in the gas outlet - C _L mole/dm³ oxygen concentration in liquid phase - C _L ^* mole/dm³ saturated oxygen concentration in liquid phase - C _0,in mole fraction of oxygen in the gas inlet - C _0,out mole fraction of oxygen in the gas outlet - CPR mole/h/dm³ carbon dioxide production rate based on total broth - E(t) error signal - F _in mole/h inlet gas flow rate - k ₁ constant defined by Eq. (4) - k ₂ constant defined by Eq. (5) - k _L a 1/h volumetric mass transfer coefficient of gas-liquid phase - OUR mole/h/dm³ oxygen uptake rate based on total broth - P atm pressure in the reactor - t h time - TS _c g total sugar consumption - TS _s g/dm³ set point of total sugar concentration - TS _* g/dm³ reference value of total sugar concentration - TS(t) g/dm³ total sugar concentration in the broth at timet - u(t) cm³/min feed rate at timet - V dm³ total broth volume - VVM (dm³/min)/dm³ flow rate per unit liquid volume - a negative constant defined by Eq. (7)     

In-situ recovery of butanol during fermentation

End-product inhibition in the acetone-butanol fermentation was reduced by using extractive fermentation to continuously remove acetone and butanol from the fermentation broth. In situ removal of inhibitory products from Clostridium acetobutylicum resulted in increased reactor productivity; volumetric butanol productivity increased from 0.58 kg/(m³h) in batch fermentation to 1.5 kg/(m³h) in fed-batch extractive fermentation using oleyl alcohol as the extraction solvent. The use of fed-batch operation allowed glucose solutions of up to 500 kg/m³ to be fermented, resulting in a 3.5- to 5-fold decrease in waste water volume. Butanol reached a concentration of 30–35 kg/m³ in the oleyl alcohol extractant at the end of fermentation, a concentration that is 2–3 times higher than is possible in regular batch or fed-batch fermentation. Butanol productivities and glucose conversions in fed-batch extractive fermentation compare favorable with continuous fermentation and in situ product removal fermentations.List of Symbols C _g kg/m³ concentration of glucose in the feed - C _w dm³/m³ concentration of water in the feed - F(t) cm³/h flowrate of feed to the fermentor at time t - V(t) dm³ broth volume at time t - V _i dm³ initial broth volume - V _si dm³ volume of the i-th aqueous phase sample - effective fraction of water in the feed Part 1. Bioprocess Engineering 2 (1987) 1–12     

Malate degradation by Schizosaccharomyces yeasts included in alginate beads

Schizosaccharomyces yeasts can be used for deacidification of grape musts. To this aim, we studied malic acid degradation by yeasts included in double layer alginate beads in a bubble column reactor. Use of immobilized micro-organisms allowed a continuous process with high dilution rates giving a deacidification capacity of 0.032 g of malate/hour/dm³/g of beads. The pneumatic agitation was very convenient in this case.List of Symbols D h^–1 Dilution rate for continuous culture - h Residence time for continuous culture - dM/dt kg/(m³ · h) Rate of degradation of malic acid - dS/dt kg/(m³ · h) Rate of consumption of glucose - _max h^–1 Maximal specific rate of growth     

Ethanol fermentation using a glucose negative mutant ofZymomonas mobilis

Summary The fermentation of an equimolar mixture of glucose and fructose into ethanol and sorbitol by a glucose negative mutant ofZymomonas mobilis was monitored. The results were analyzed using a recently described method based on polynomial fitting and calculation of intantaneous and overall parameters. These parameters described well the physiology of this mixed-substrate mixed-product fermentation. Growth of the mutant was greatly inhibited on this medium. Fructose was quantitatively converted into sorbitol while glucose was oxidized into gluconic acid .This latter product was utilized as substrate for cell growth and ethanol production.Nomenclature X biomass concentration, g/l - S total sugar concentration, g/l - Glu glucose concentration, g/l - Fru fructose concentration, g/l - Sor sorbitol concentration, g/l - P ethanol concentration, g/l - t fermentation time, h - specific growth rate, h^-1 - q_s specific sugar uptake rate, g/g.h - q_G specific glucose uptake rate, g/g.h - q_F specific fructose uptake rate, g/g.h - q_P specific ethanol productivity, g/g.h - q_Sor specific sorbitol productivity, g/g.h - Y_X/S biomass yield on total sugar, g/g - Y_P/S ethanol yield on total sugar, g/g - Y_Sor/S sorbitol yield on total sugar, g/g - y_Sor/f sorbitol yield on fructose, g/g - Y_P/G ethanol yield on glucose, g/g     

Procedure for determination of elemental sulphur bio-oxidation rate

A procedure is described for measuring the rate of biooxidation of elemental sulphur in nutrient solutions. Results of preliminary measurements of sulphur bio-oxidation rate in a dynamic system are presented. The rate of sulphur bio-oxidation has been determined at the level of 0.02–0.05 g of sulphur per m² of sulphur per h.List of Symbols C g/dm³ concentration of sulphate ions - C ₂ g/dm³ concentration of sulphate ions in withdrawn solution - C g/dm³ C difference between solution outlet and inlet to sulphur bed - F m² sulphur surface exposed to bacteria action - m g mass of elemental sulphur - V dm³ volume of solution - V ₀ dm³/h volume of fresh solution supplied to the set - V ₁ dm³/h circulating solution flow rate - V ₂ dm³/h volume of solution withdrawn - h time Abbreviations RBES rate of bio-oxidation of elemental sulphur     

Process engineering aspects of the bioleaching of copper ores

The bioleaching of minerals is a complex process that is affected by a number of biological, mineralogical, electrochemical and engineering factors. This work presents and discusses the most significant process engineering aspects involved in the bacterial leaching of copper ores, i.e. bacterial population, type of mineral and particle size, nutrients and inhibitors, oxygen and carbon dioxide, temperature and pH, leaching kinetics and operation mode.It is concluded that more work is needed in this area in order to gain a deeper insight in the many factors that govern this process. This would allow to significantly improve its overall productivity.List of Symbols C _L kg/m³ dissolved oxygen concentration - C ^* kg/m³ equilibrium oxygen concentration - d, e, f, g % percentage of C, H, O and N in the cell - D m impeller diameter - K consistency index - K _S, K₁, K_c constants - k _La h^–1 volumetric oxygen transfer coefficient - M _b mol/kg biomass apparent molecular weight - N s^–1 rotation frequency - n behavior index - P kg/m³ ungassed agitation power, product concentration - P _g kW/m³ gassed agitation power - p % pulp density - Q m³/h air flow rate - S kg/m³ limiting substrate concentration - W kg/(m³ · h) mass transfer rate per unit volume - X cells/cm³ biomass concentration - Y _o g cells/g Fe oxygen cell yield - Y _x g cells/g Fe substrate cell yield - h^–1 specific growth rate - _m h^–1 maximum specific growth rate     

Theoretical analysis of the baker's yeast production: An experimental verification at a laboratory scale

The scale-down procedure can be used to optimize and scale up fermentation processes. The first step in this procedure, a theoretical analysis of the process at a large scale, must give information about the regime, or bottle necks, ruling the process. In order to verify the theoretical results the process analysis has been applied to the fed-batch baker's yeast production at a laboratory scale. The results of this analysis are compared with results from fed-batch experiments. It was concluded that if only one mechanism is ruling the process, for instance mass transfer, the results of the analysis are quite clear. If more than one mechanism is important, for example mass transfer and liquid mixing, additional knowledge is needed to predict the behaviour of the process.Concerning the baker's yeast production, it was concluded that if oxygen limitation occurs, liquid mixing is of little importance.List of Symbols C kg/m³ concentration - C ^* kg/m³ saturation concentration - D m diameter - D _E m²/s effective dispersion coefficient - d m holes of the sparger - F _sm³/s substrate flow to the fermentor - g m/s² gravitational acceleration - H m height - k _La s^–1 volumetric mass transfer coefficient based on the liquid volume - L m length - m _skg/(kg·s) maintenance coefficient - OTR kg/(m³·s) oxygen transfer rate - OUR kg/(m³·s) oxygen uptake rate - r kg/(m³·s) reaction rate - t s time - V m³ volume - v m/s velocity - v _sm/s superficial gas flow rate - y _ijkg/kg yield of componentj oni - s^–1 specific growth rate - s time constant - _gm³/s gas flow rate Indices 0 value att=0 - cir liquid circulation - e ethanol - f feed concentration - g gas phase - in flow going to the fermentor - l liquid phase - m mixing - mt mass transfer - o, O₂ oxygen - oc oxygen consumption - out flow coming out the fermentor - s substrate - sa substrate addition - sc substrate consumption - x biomass     

Ethanol fermentation using a fructose-negative mutant of Zymomonas mobilis

Summary The fermentation of an equimolar mixture of glucose and fructose into ethanol and sorbitol by a fructose negative mutant of Zymomonas mobilis is analysed using a recently described methodology (Ait-Abdelkader and Baratti, Biotechnol. Tech. 1993,329–334) based on polynomial fitting and calculation of instantaneous and overall parameters. These parameters are utilized to describe this mixed-substrate mixed-product fermentation.Nomenclature X biomass concentration, g/l - S total sugar concentration, g/l - Glu glucose concentration, g/l - Fru fructose concentration, g/l - Sor sorbitol concentration, g/l - P ethanol concentration, g/l - t fermentation time, h - specific growth rate, h^-1 - q_s specific sugar uptake rate, g/g.h - q_g specific glucose uptake rate, g/g.h - q_F specific fructose uptake rate, g/g.h - q_P specific ethanol productivity, g/g.h - q_Sor specific sorbitol productivity, g/g.h - Y_X/S biomass yield on total sugar, g/g - Y_P/S ethanol yield on total sugar, g/g - Y_Sor/S sorbitol yield on total sugar, g/g - Y_Sor/F sorbitol yield on fructose, (g/g) - Y_P/G ethanol yield on glucose, (g/g)     

Phenol degradation in a three-phase biofilm fluidized sand bed reactor

A previous three phase fluidized sand bed reactor design was improved by adding a draft tube to improve fluidization and submerged effluent tubes for sand separation. The changes had little influence on the oxygen transfer coefficients(K _L a), but greatly reduced the aeration rate required for sand suspension. The resulting 12.5 dm³ reactor was operated with 1 h liquid residence time, 10.2dm³/min aeration rate, and 1.7–2.3 kg sand (0.25–0.35 mm diameter) for the degradation of phenol as sole carbon source. The K _La of 0.015 s^–1 gave more than adequate oxygen transfer to support rates of 180g phenol/h · m³ and 216 g oxygen/h · m³. The biomass-sand ratios of 20–35 mg volatiles/g gave estimated biomass concentrations of 3–6 g volatiles/dm³. Offline kinetic measurements showed weak inhibition kinetics with constants ofK _s=0.2 mg phenol/dm³, K _o2=0.5 mg oxygen/dm³ and KinI= 122.5 mg phenol/dm³. Very small biofilm diffusion effects were observed. Dynamic experiments demonstrated rapid response of dissolved oxygen to phenol changes below the inhibition level. Experimentally simulated continuous stagewise operation required three stages, each with 1 h residence time, for complete degradation of 300 mg phenol/dm³ · h.     

Studies on RTD and continuous culture (SCP) in cylindrical and tapered reactors

The performance of a tapered reactor for the continuous cultivation of bakers' yeast (SCP) from cane molasses has been compared with that of a conventional cylindrical reactor. It is found that the tapered reactor has less non-idealities (bypass and deadspace).Using the experimentally evaluated bypass and deadspace values, a model for predicting conversions of substrate (cane molasses), based on the RTD model proposed by Cholette and Cloutier has been developed. The experimental substrate conversions are found to match the model satisfactorily.List of Symbols D h^–1 dilution rate - E() exit age distribution function - K _s kg/m³ Monod's saturation constant - -r _sa kg/(m³ · h) rate of substrate utilization - S kg/m³ substrate concentration expressed as dextrose equivalent (DE) - S _a kg/m³ substrate concentration in active zone - S ₀ kg/m³ initial substrate concentration - S/S ₀ dimensionless substrate concentration - v _a dm³/h volumetric flow through active zone - v _b dm³/h volumetric flow through bypass stream - u _l dm³/h substrate feed rate - v _g dm³/min air-flow rate - V dm³ total working volume of the reactor - V _a dm³ volume of active zone in reactor - V _d dm³ volume of dead zone in reactor - X kg/m³ biomass concentration Greek Letters fraction of bypass of feed, v _b /v _l - fraction of deadspace, V _d /V - dimensionless residence time - _m h^–1 maximum specific growth rate - h mean residence time, V/v _l      

Mathematical modeling of production of microbial lipids

As a part of the investigations on the microbial lipid production using the yeast Rhodotorula gracilis, CFR-1, kinetics of the biomass synthesis has been studied using shake flask experiments. Using a medium containing a carbon to nitrogen ratio of 701, the rates of biomass production were followed at different initial substrate concentrations in the range of 20–100 kg/m³. A logistic model was found to be reasonably adequate to describe the kinetics of the growth of biomass; the maximum specific growth rate of 0.105 h^–1 was applicable for substrate concentrations less than 60 kg/m³, which gave reasonable agreement between predicted and actual biomass concentration values.List of Symbols S ₀, X ₀ kg/m³ Initial concentrations of sugar, non lipid biomass respectively - X, X(t) kg/m³ Concentrations of non lipid biomass at any time t - dX/dt kg/(m³ · h) Rate of biomass growth - h^–1 Specific growth rate - _max h^–1 Maximum specific growth rate - K _s mol/dm³ Monods constant - X _max kg/m³ Maximum biomass reached in a run     

Growth of E. coli in a stirred tank and in an air lift tower reactor with an outer loop

E. coli ATCC 11105 was cultivated in a 10-1 stirred tank reactor and in a 60-1 tower loop reactor in batch and continuous operation. By on-line measurements of O₂ and CO₂ concentrations in the outlet gas, pH, temperature, cell mass concentration X as well as dissolved O₂ concentration along the tower in the broth, gas holdup, broth recirculation rate through the loop and by offline measurements of substrate concentration DOC and cell mass concentration along the tower, the maximum specific growth rate _m , yield coefficients Y _X/S. Y _X/DOC and were evaluated in stirred tank and tower loop in batch and continuous cultures with and without motionless mixers in the tower and at different broth circulation rates through the loop. To control the accuracy of the measurements the C balance was calculated and 95% of the C content was covered.The biological parameters determined depend on the mode of operation as well as on the reactor used. Furthermore, they depend on the recirculation rate of the broth and built-ins in the tower. The unstructured cell and reactor models are unable to explain these differences. Obviously, structured cell and reactor models are needed. The cell mass concentration can be determined on line by NADH fluorescence in balanced growth, if the model parameters are determined under the same operational conditions in the same reactor.List of Symbols a, b empirical parameters in Eq. (1) - CPR kg/(m³ h) CO₂ production rate - C kg/m³ concentration - D l/h dilution rate - DOC kg/m³ dissolved organic carbon - I net. fluorescence intensity - K _S kg/m³ Monod constant - k _L a l/h volumetric mass transfer coefficient - OTR kg/(m³ h) oxygen transfer rate - OUR kg/(m³ h) oxygen utilization rate - RQ = CPR/OUR respiratory quotient - S kg/m³ substrate concentration - t h,min, s time - t _u min recirculation time - t _M min mixing time - v m³/h volumetric flow rate through the loop - X kg/m³ (dry) cell mass concentration - Y _X/S yield coefficient of cell mass with regard to the consumed substrate - Y _X/DOC yield coefficient of the cell mass with regard to the consumed DOC - Y _X/O yield coefficient of the cell mass with regard to the consumed oxygen - Z relative distance in the tower from the aerator with regard to the height of the aerated broth - l/h specific growth rate - _m l/h maximum specific growth rate Indices f feed - e outlet     

Investigations of cephalosporin C production in an airlift tower loop reactor

Summary Cephalosporin C was produced by Cephalosporium acremonium in a 60 l airlift loop reactor on complex medium (with 30 kg/m³ peanut flour) in fed-batch operation. A final product concentration of 5 kg/m³ and a maximum productivity of 45 g/m³ h were attained. On-line analysis was used to determine ammonia, methionine, phosphate, reducing sugar and cephalosporin C by an autoanalyser, glucose by a flow injection analyser and cephalosporin C, penicillin N, deacetoxycephalosporin C, deacetylce-phalosporin C and methionine by HPLC. The volumetric productivity of the stirred tank reactor was higher than that of the airlift reactor because of differences in cell concentration. Specific productivities in relative to cell mass were similar in the two reactors. The substrate yield coefficient in the airlift reactor was twice that in the stirred tank reactor.Nomenclature E _o2 efficiency of oxygen transfer with regard to the specific power input - K _La volumetric mass transfer coefficient - OTR oxygen transfer rate - P power input - PR volumetric productivity of CPC - q _a volumetric aeration rate/broth volume (vvm) - SPR specific productivity with regard to RNA - V _L broth volume in reactor - z relative height of the aerated reactor     

Biotransformation of cephalosporin C to 7-aminocephalosporanic acid with coimmobilized biocatalyst in a batch bioreactor

The permeabilized cells of Trigonopsis variabilis CCY 15-1-3 having D-amino acid oxidase (DAAO) activity were used to convert cephalosporin C (CPS-C) into 7-(-ketoadipyl amido) cephalosporanic acid (CO-GL-7-ACA) in a batch bioreactor with good aeration and stirring during the process. The deacylation of 7--(4-carboxybutanamido)-cephalosporanic acid (GL-7-ACA) to 7-cephalosporanic acid (7-ACA) by permeabilized cells of Pseudomonas species 3635 having 4--(4-carboxybutamido)-cephalosporanic acid acylase (GL-7-ACA acylase) activity was performed in a batch bioreactor. A spectrophotometric method for the determination of CO-GL-7-ACA and 7-ACA was proposed. Experimental data were fitted by non-linear regression with parameters optimization. The sorption method (without reaction) was applied for the determination of cephalosporin effective diffusion coefficients in Ca-pectate gel beads. These beads were prepared by dropping a potassium pectate gel suspension of inactive permeabilized cells of Trigonopsis variabilis and Pseudomonas species, crosslinked with glutaraldehyde, into a stirred 0.2 M calcium chloride solution. Concentrations of appropriate cephem components were measured by the refractive method. Values of effective diffusion coefficients were calculated by the Fibbonacci optimization method.List of Symbols c _L mol/dm³ concentration on the surface of a bead - c _L0 mol/dm³ initial cephalosporin concentration - c _L mol/dm³ equilibrium cephalosporin concentration in the solution - c _s1 mol/dm³ concentration of CPS-C - c _s2 mol/dm³ concentration of GL-7-ACA - D _ei m²/s effective diffusion coefficient of the components - K _i mol/dm³ inhibition parameter in Eq. (2) - K _{m i} mol/dm³ Michaelis constant in Eq. (1) - K _{m 2} mol/dm³ Michaelis constant in Eq. (2) - n number of beads - q _n nonzero positive roots in Eq. (7) - r ₁ mol/(dm³·s) rate of the conversion of CPS-S to CO-GL-7-ACA - r ₂ mol/(dm³·s) rate of the conversion of GL-7-ACA to 7-ACA - R m radius of the bead - S( ) symbol for total residual sum of squares in Eq. (1) - t s time - V _{m 1} mol/(dm³·s) max. reaction rate in Eq. (1) - V _{m 2} mol/(dm³·s) max. reaction rate in Eq. (2) - V _L dm³ volume of the solution excluding the space occupied by beads - V _s dm³ volume of beads - y _i mol/(dm³ · s) symbol for experimental data in Eq. (1) - _i mol/(dm³· s) symbol for calculated data in Eq. (1) - _P porosity, defined by Eq. (5) - dimensionless parameter, defined by Eq. (6) The authors wish to thank Dr. P. Gemeiner of Slovak Academy of Sciences for rendering of pectate gel. This work is supported by Ministry of Education (Grant No. 1/990 935/93)