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     

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     

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      

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      

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     

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     

Simulation of the dynamics in the Baker's yeast process

Simulation of the dynamics in a fed batch process for production of Baker's yeast is discussed and applied. Experimental evidences are presented for a model of the energy metabolism. The model involves the concept of a maximum respiratory capacity of the cell. If the sugar concentration is increased above a critical value, corresponding to a critical rate of glycolysis and a maximum rate of respiration, then all additional sugar consumed at higher sugar concentrations is converted into ethanol.In a fed batch process with constant sugar feed the sugar concentration declines slowly. If ethanol is present when the sugar concentration declines below the critical value of 110 mg/dm³ fructose +glucose the metabolism switches rapidly into combined oxidation of sugar and ethanol. Thus, no diauxic growth is involved under process conditions. The rate of ethanol consumption is determined by the free capacity of respiration under these conditions. The invertase activity of the cells was found to be so high that mainly fructose and glucose were present in the medium, typically in the concentration range around 100 mg/dm³. These components are consumed at the same rate but with fructose at a higher concentration, indicating a higher K _s for fructose consumption.The model was used in simulation experiments to demonstrate the dynamics of the Baker's yeast process and the influence of different process conditions.List of Symbols DOT % air sat dissolved oxygen tension - F dm³/h rate of inlet medium flow - H kg/(dm³ % air sat.) oxygen solubility - K kg/m³ saturation constant specified by index - K _L a 1/h volumetric oxygen transfer coefficient - m g/(g · h) maintenance coefficient specified by index - P kg/(m³ · h) mean productivity of biomass in the process - q g/(g · h) specific consumption or production rate - S kg/m³ concentration of sugar in reactor - S ₀ kg/m³ concentration of inlet medium sugar medium t h process time - V dm³ medium volume - X kg/m³ concentration of biomass - Y g/g yield coefficient specified by index - 1/h specific growth rate Index aa anaerobic condition - c critical value - e ethanol - ec ethanol consumption - ep ethanol production - max maximum value - o oxygen - oe oxygen for growth on ethanol - os oxygen for growth on sugar - s sugar - x biomass     

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     

Mathematical models for mixing in deep-jet bioreactors: Calculation of parameters

The macroscopic mathematical model based on compartments with ideal mixing zones and tanks-in series was evaluated. Based on the experimental data obtained in a 300 dm³ pilot reactor and the dependence of mixing time on the volume of liquid phase, we have found mathematical relations between the ratio of vessel diameter to liquid level, adjustable parameters of model and the mixing time.List of Symbols V dm³ total volume of bioreactor - V _g dm³ total volume of liquid - V ₁ dm³ volume of ideally mixed zone in the vessel - V ₂ dm³ volume of macromixer in inner circulation flows - V ₃ dm³ volume of liquid phase in the pump - V ₄ dm³ volume of liquid phase in the pipe between the vessel and the pump - V ₅ dm³ volume of liquid phase in the pipe between the pump and air input system included falling jet - V _LT dm³ volume of liquid in the tank - V _LC dm³ volume of liquid in the circulation system - F _E dm³/s inner volumetric circulation flow rate across the macromixers - F _cir dm³/s external volumetric circulation flow rate, pumping capacity - t _A s time interval of the pulse application - t _AA s time point of the pulse application related to the free choosen starting point of the experiment - t _m s mixing time - t _c s circulation time - t _end s end time of simulation - C _*,* kg/m³ concentration of tracer in the indicated compartment - C ₀ kg/m³ concentration of the tracer before the injection - C _t kg/m³ concentration of the tracer at the indicated time - C kg/m³ theoretical concentration of the full mixed tracer - C _sim kg/m³ calculated concentration of tracer during numerical integration method - i index of an arbitrary tank - D _T m diameter of bioreactor - D 1/s dilution rate - H _L m level of liquid in the unaerated vessel - vector of inhomogenities     

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     

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)     

Bioprocess engineering characteristics of the horizontal stirred tank

Bioreactor performance studies of the recently developed horizontal stirred tank with a volume of 421 have been carried out for fermentation with Trichosporon cutaneum. Quantification on the basis of measured oxygen transfer capacity and power consumption is presented and compared with data for a conventional vertical tank bioreactor.During the experiments it has been observed that two different forms of morphology of Trichosporon, i.e. the normal yeast-form (Y) with single cells and a mycelium-form (M) with filamentous cells, are present in the horizontal stirred tank when working with the original strain (DSM 70698). After separation both forms were characterized and later on used for bioreactor performance studies in the horizontal and vertical stirred tank. Results of oxygen efficiency show the drastic effect of the morphology change on bioreactor performance. Finally different bioreactors are quantitatively compared on the basis of oxygen transfer, power consumption and productivity using the reference fermentation system Trichosporon cutaneum.List of Symbols F m³/h flow rate (volumetric) - k _La1/h volumetric transfer coefficient of OTR - M Nm torque - n 1/s rotational speed - P Nm/s power - V m³ volume - V _G1/min gas flow rate - x kg/m³ biomass concentration - ^* morphology index - ^* engineering (specific) viscosity - _app Ns/m² apparent viscosity - ₀ N/m² yield stress (Casson law) - t _1/e h measured time acc. to momentum method [17] - tEh characteristic time of electrode response - t _Gh mean residence time of gas phase Abbreviations CFR completely filled reactor - CRR cyclic ring reactor (torus) - JLR jet loop reactor - HSTR horizontal stirred tank reactor - M mycelium-form of Trichosporon cutaneum - O₂-eff O₂-efficiency - OUR O₂-uptake rate - OTR O₂-transfer rate - STR stirred tank reactor - ThLR thin layer reactor - VSTR vertical stirred tank reactor - Y yeast-form of Trichosporon cutaneum The work presented in this paper was supported by an Austrian Research Grant (FFWF, Project no. 4496)     

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     

Performance of cell immobilized packed bed bioreactor for continuous fermentation of alcohol

Experiments were conducted in a packed bed bio-reactor consisting of entrapped yeast cells in alginate matrix for continuous production of alcohol. The variables include initial substrate level, reactor diameter, diameter of the bead and residence time. The influence of these parameters on the conversion of substrate was studied. The film and pore diffusional effects were observed by varying the column and bead diameters, respectively. The pseudo first order reaction rate constant was calculated and correlated with the bead diameter. The effectiveness factor and the Thiele modulus were estimated. A correlation was proposed for fractional conversion in terms of operating variables. It is possible to predict the residence time required and volumetric productivity achieved in a bioreactor for any given initial substrate concentration at any fractional conversion obtained.List of Symbols a _m m²/kg surface are per unit mass of catalyst particle - D m diameter of the reactor - D _e m²/s effective diffusivity - d m particle diameter - h m bed height - k m/s first order reaction rate constant - k m³/(kg · s) pseudo first order reaction rate constant - k _in m³/(kg · s) intrinsic reaction rate constant, (=K/gh) - k _m m/s mass transfer coefficient - P kmol/(m³ · s) volumetric productivity - Q m³/s flow rate of the feed - S kmol/m³ substrate concentration at any time - S _o kmol/m³ initial substrate concentration - S _p kmol/m³ substrate concentration on the gel bead surface - t s reaction time - T (kg · cat · s)/m³ space time (weight of the biocatalyst/flow rate of the feed) - v kmol/(kg · cat · s) reaction rate - V _pfr m³ volume of the packed bed reactor - X [1-(S/S _o)] fraction of the substrate converted in to product Greek Symbols effectiveness factor - Thiele modulus - kg/m³ density of the catalyst particle - s residence time, (= D² h/4Q) - voidage     

Modelling of alcohol fermentation in a tubular reactor with high biomass recycle

Fermentation in tubular recycle reactors with high biomass concentrations is a way to boost productivity in alcohol production. A computer model has been developed to investigate the potential as well as to establish the limits of this process from a chemical engineering point of view. The model takes into account the kinetics of the reaction, the nonideality of flow and the segregation in the bioreactor. In accordance with literature, it is shown that tubular reactors with biomass recycle can improve productivity of alcohol fermentation substantially.With the help of the computer based reactor model it was also possible to estimate the detrimental effects of cell damage due to pumping. These effects are shown to play a major role, if the biomass separation is performed by filtration units which need high flow rates, e.g. tangential flow filters.List of Symbols Bo d Bodenstein number - c kg/m³ concentration of any component - CPFR continuous plug flow reactor - CSTR continuous stirred tank reactor - d _h m hydraulic diameter - D _eff m²/s dispersion coefficient - f residence time distribution function - K _s kg/m³ monod constant for biomass production - K _s kg/m³ monod constant for alcohol production - p kg/m³ product concentration - P _i kg/m³ lower inhibition limit concentration for biomass production - p _i kg/m³ lower inhibition limit concentration for alcohol production - p _m kg/m³ maximum inhibition limit concentration for biomass production - p _m kg/m³ maximum inhibition limit concentration for alcohol production - q _p h^–1 specific production rate - q _p,max h^–1 maximum specific production rate for alcohol production - q _s h^–1 specific substrate consumption rate - Q _L m _gas ³ /m³h specific gas rate - r _p , r _s , r _x kg/(m³ · h) reaction rate for ethanol production substrate consumption and cell growth, respectively - S _F kg/m³ substrate concentration in feed stream - s kg/m³ substrate concentration - t h time - x kg/m³ biomass concentration - x _max kg/m³ maximum biomass concentration for biomass production - Y _p/s yield coefficient - h^–1 specific growth rate - _max h^–1 maximum specific growth rate - dimensionless time (t/) - h mean residence time - _s glucose conversion     

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     

Optimization of the microbial synthesis of dihydroxyacetone from glycerol with <Emphasis Type="Italic">Gluconobacter oxydans</Emphasis>

An optimized repeated-fed-batch fermentation process for the synthesis of dihydroxyacetone (DHA) from glycerol utilizing Gluconobacter oxydans is presented. Cleaning, sterilization, and inoculation procedures could be reduced significantly compared to the conventional fed-batch process. A stringent requirement was that the product concentration was kept below a critical threshold level at all times in order to avoid irreversible product inhibition of the cells. On the basis of experimentally validated model calculations, a threshold value of about 60 kg m^-3 DHA was obtained. The innovative bioreactor system consisted of a stirred tank reactor combined with a packed trickle-bed column. In the packed column, active cells could be retained by in situ immobilization on a hydrophilized Ralu-ring carrier material. Within 17 days, the productivity of the process could be increased by 75% to about 2.8 kg m^-3 h^-1. However, it was observed that the maximum achievable productivity had not been reached yet.Abbreviations K _O Monod half saturation constant of dissolved oxygen (kg m^-3) - K _S Monod half saturation constant of substrate glycerol (kg m^-3) - O Dissolved oxygen concentration (kg m^-3) - P Product concentration (kg m^-3) - P _crit Critical product concentration constant (kg m^-3) - S Substrate concentration (kg m^-3) - t Time (s) - X Biomass concentration (dry weight) (kg m^-3) - Y _P/S Yield coefficient of product from substrate - Y _X/S Yield coefficient of biomass from substrate - Growth dependent specific production rate constant (kg m^-3) - Growth independent specific production rate constant (s^-1) - Specific growth rate (s^-1) - _max Maximum specific growth rate constant (s^-1)     

The effect of micro- and macromixing on enzyme reaction in a real CSTR

The effect of micromixing and macromixing on enzyme reaction of Michaelis-Menten type in a real continuously stirred tank reactor (CSTR) is considered. The effect of bypassing of a fraction of feed stream, dead space, initial enzyme concentration and Michaelis-Menten constant on substrate conversion is evaluated. Bypass reduces the substrate conversion significantly compared with other parameters in the case of micro and macromixing. Micromixing predicts higher substrate conversions compared with macromixing. The effect of micro and macromixing on substrate conversion is negligible at low and high conversions.List of Symbols C kmol/m³ concentration of reactant - ¯C kmol/m³ average concentration of reactant - C_A kmol/m³ exit concentration of reactant A - C_Aa kmol/m³ exit concentration of reactant A from active zone - C_AO kmol/m³ initial concentration of reactant A - C_EO kmol/m³ initial enzyme concentration - C_O kmol/m³ initial concentration of reactant - E(t) 1/s exit age distribution function - k 1/s reaction rate constant - M kmol/m³ Michaelis-Menten constant - r kmol/(m³s) rate of reaction - –r_A kmol/(m³s) rate of reaction with respect to A - t s time - v m³/s volumetric feed rate - v_a m³/s volumetric feed rate entering the active zone - v_b m³/s volumetric feed rate entering the bypass stream - V m³ total volume of the vessel - V_a m³ active volume of the vessel - V_d m³ volume of dead space - X_A conversion of A Greek Letters fraction of feed stream bypassing the vessel (v_b/v) - fraction of the total volume as dead space (V_d/V) - (t) 1/s Dirac delta function, an ideal pulse occurring at time t = 0 - s life expectancy of a molecule - 1/s intensity function or escape probability function - s space time or mean residence time     

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

The scale-down procedure seems an adequate tool in the design, optimization and scale-up fermentation processes. The first step in this procedure is a theoretical analysis, called process analysis, which is based on characteristic times of the mechanisms which may influence the performance of the bioreactor. This analysis must give information about the behaviour of large and small scale fermentation processes. At a small scale a verification of the results of such an analysis of the fed-batch baker's yeast production is carried out.In this paper a comparison of calculated and measured characteristic times of liquid mixing and mass transfer is presented. It was concluded that the literature correlations give a rough estimation of the characteristic times and can be used in the process analysis. Depending on the kind of sparger, the medium and the scale of the reactor, more knowledge is needed about bubble coalescence in fermentation media.The volumetric oxygen transfer coefficient increased when the biomass concentration increased. Probably this is caused by the interaction between biomass and the anti-foaming agent used.List of Symbols C kg/m³ concentration - D m diameter - m²/s effective dispersion coefficient - d m holes of the sparger - g m/s² gravitational acceleration - H m height - k _L a s^–1 volumetric mass transfer coefficient based on the liquid volume - L m length - m kg/kg gas liquid distribution coefficient - OTR kg/(m³ · s) oxygen transfer rate - OUR kg/(m³ · s) oxygen uptake rate - t s time - ^s m/s superficial gas flow rate - m length - s time constant - _g m³/s gas flow rate Indices 0 value at t=0 - cal calculated - e value at t=t (end) - g gas phase - in flow going to the fermentor - l liquid phase - m mixing - mt mass transfer - O ₂ oxygen - out flow coming out the fermentor     

Reactor design for the enzymatic isomerization of glucose to fructose

A comprehensive methodology is presented for the design of reactors using immobilized enzymes as catalysts. The design is based on material balances and rate equations for enzyme action and decay and considers the effect of mass transfer limitations on the expression of enzyme activity. The enzymatic isomerization of glucose into fructose with a commercial immobilized glucose isomerase was selected as a case study. Results obtained are consistent with data obtained from existing high-fructose syrup plants. The methodology may be extended to other cases, provided sound expressions for enzyme action and decay are available and a simple flow pattern within the reactor might be assumed.List of Symbols C kat/kg specific activity of the catalyst - D m²/s substrate diffusivity within the catalyst particle - Dr m reactor diameter - d d operating time of each reactor - E kat initial enzyme activity - E _i kat initial enzyme activity in each reactor - F m³/s process flowrate - F _i m³/s reactor feed flowrate at a given time - F ₀ m³/s initial feed flowrate to each reactor - H number of enzyme half-lives used in the reactors - K mole/m³ equilibrium constant - K _S mole/m³ Michaelis constant for substrate - K _P mole/m³ Michaelis constant for product - K _m mole/m³ apparent Michaelis constant f(K, K _s, K_p, s₀) - k mole/s · kat reaction rate constant - k _d d^–1 first-order thermal inactivation rate constant - L m reactor height - L _r m height of catalyst bed - N _R number of reactors - P _i kg catalyst weight in each reactor - p mole/m³ product concentration - R m particle radius - R _P ratio of minimum to maximum process flowrate - r m distance to the center of the spherical particle - s mole/m³ substrate concentration - s _0i mole/m³ substrate concentration at reactor inlet - s ₀ mole/m³ bulk substrate concentration - s mole/m³ apparent substrate concentration - T K temperature - t d time - t _i d operating time for reactor i - t _s d time elapsed between two successive charges of each reactor - V m³ reactor volumen - V _m mole/m³ s maximum apparent reaction rate - V _p mole/m³ s maximum reaction rate for product - V _R m³ actual volume of catalyst bed - V _r m³ calculated volume of catalyst bed - V _S mol/m³ s maximum reaction rate for substrate - v mol/m³ s initial reaction rate - v _i m/s linear velocity - v _m mol/m³ s apparent initial reaction rate f(K_m, s,V_m) - X substrate conversion - X _eq substrate conversion at equilibrium - =s/K dimensionless substrate concentration - ₀=s₀/K bulk dimensionless substrate concentration - _eq=s_eq/K dimensionless substrate concentration at equilibrium - local effectiveness factor - mean integrated effectiveness factor - Thiéle modulus - =r/R dimensionless radius - _s kg/m³ hydrated support density - substrate protection factor - s residence time