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     

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     

Cephalosporin C production in a stirred tank reactor

Summary Cephalosporin C was produced with the moldCephalosporium acremonium in a 20 1 stirred tank reactor with 100 kg/m³ peanut flour in fed-batch operation. The growth and product formation was followed by on-line analysis of the broth composition. The cell concentration was estimated from the RNA-content of the cells. By optimization of the fed-batch operation and by increasing the phosphate content in the broth, a final cephalosporin C concentration of 12 kg/m³ was attained.Nomenclature CPC cephalosphorin C - DAC deacetylcephalosporin C - DAOC deacetoxycephalosporin C - k _L a volumetric mass transfer coefficient - MMBS 2-Hydroxy-4-methylmercaptobutyric acid - PABAH p-Hydroxybenzoicacidhydrazid - RNA ribonucleic acid - RQ respiratory quotient - oxygen transfer rate - CO₂-production rate - t fermentation time     

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     

Model for the production of L-tryptophan from L-serine and indole by immobilized cells in a three-phase liquid-impelled loop reactor

Production of L-tryptophan from L-serine and indole catalyzed by Escherichia coli, immobilized in k-carrageenan gel beads, is technically feasible in the liquidimpelled loop reactor (LLR), using an organic solvent, e.g. n-dodecane.With L-serine in large excess intrinsic reaction kinetics is approximately first order with respect to indole, with a reaction constant of 8.5×10^–5 m³ kg _dw ^–1 s^–1.The overall process kinetics is jointly controlled by intrinsic kinetics and by intraparticle mass transfer resistance, which can be quantified using an effectiveness factor.Mass transfer of indole from the organic to the aqueous phase and from the aqueous to the gel phase are relatively fast and thus have negligible influence in the overall process kinetics, under the operational conditions tested. However, they may become important if the process is intensified by increasing the cell concentration in the gel and/or the gel hold-up in the reactor.A simple model which includes indole mass balances over the aqueous and organic phases, mass transfer and reaction kinetics, with parameters experimentally determined in independent experiments, was successful in simulating L-tryptophan production in the LLR.List of Symbols a, b, c coefficients of the equilibrium curve for indole between organic and aqueous phases - A, B, C, D, E, F auxiliary variables used in liquid-liquid mass transfer studies - a _x specific interfacial area referred to the volume of the aqueous phase (m^–1) - A _x interfacial area (m²) - a _Y specific interfacial area referred to the volume of the organic phase (m^–1) - A _Y interfacial area (m²) - C _b substrate concentration in the bulk of the aqueous phase (kg m^–3) - C _e substrate concentration in exit stream (kg m^–3) - C _E biocatalyst concentration referred to the aqueous phase (kg m^–3) - C _{E s} biocatalyst concentration referred to the volume of gel (kg m^–3) - C _s substrate concentration at the gel surface (kgm^–3) - d, e, f coefficients of the equilibrium curve for indole between aqueous and organic phases - dp particle diameter (m) - K ₂ kinetic constant (s^–1) - K ₁ kinetic constant K₂/K_M (kg^–1 m³ s^–1) - K _M Michaälis-Menten constant (kgm^–3) - K _X mass transfer coefficient referred to the aqueous phase (ms^–1) - K _Xa_X volumetric mass transfer coefficient based on the volume of the aqueous phase (s^–1) - k _Y mass transfer coefficient referred to the organic phase (ms^–1) - K _Ya_Y volumetric mass transfer coefficient based on the volume of the organic phase (s^–1) - N _X mass flux of indole from organic to aqueous Phase (kg m^–2s^–1) - N _Y mass flux of indole from aqueous to organic phase (kg m^–2s^–1) - Q _e volumetric flow rate in exit stream (m³s^–1) - Q _f volumetric flow rate in feed stream (m³s^–1) - _obs observed reaction rate (kg s^–1 m^–3) - intrinsic reaction rate (kg s^–1 m^–3) - Re Reynolds number - Sc Schmidt number - Sh Sherwood number - t time (s) - u superficial velocity (m s^–1) - V _max maximum reaction rate (kg s^–1m^–3) - V _S volume of the support (m³) - V _X volume of aqueous phase (m³) - V _Y volume of the organic phase (m³) - X indole concentration in the aqueous phase (kgm^–3) - Y indole concentration in the organic phase (kg m^–3 Greek Letters overall effectiveness factor - _e external effectiveness factor - _i internal effectiveness factor - Thiele module A fellowship awarded to one of us (D.M.R.)by INICT is gratefuly acknowledged.     

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     

Fundamental aspects of waste sewage sludge treatment: Microbial solids biodegradation in an aerobic thermophilic semi-continuous system

The solubilization and biodegradation of whole microbial cells by an aerobic thermophilic microbial population was investigated over a 72 h period. Various parameters were followed including total suspended solids reduction, changes in the dissolved organic carbon, protein and carbohydrate concentrations, and carboxylic acid production and utilisation. From the rates of removal of the various fractions a simple model for the biodegradation processes is proposed and verified with respect to acetic acid production and utilization, and total suspended solids removal. The process is initiated by enzymic degradation of the substrate microbe cell walls followed by growth on the released soluble substrates at low dissolved oxygen concentration with concommitant carboxylic acid production. Subsequent utilization of the unbranched, lower molecular weight carboxylic acids allows additional energy supply following exhaustion of the easily utilisable soluble substrate from microbial cell hydrolysis.List of Symbols Y _Xp/Xs kg/kg yield process microbes on substrate yeast cells - Y _Xp/Ac kg/kg yield process microbes on acetate - Y _Ac/Ss kg/kg yield acetate produced by process microbes growing on substrate yeast cells - Y _Ss/Xs kg/kg yield soluble substrate from lysis of yeast cells - Y _Ss/Xp kg/kg yield soluble substrate from lysis of process microbes - Y _P/Xs kg/kg yield particulates from lysis of yeast cells - Y _P/Xp kg/kg yield particulates from lysis of process microbes - _{max (Ss)} h^–1 maximum specific growth rate constant for growth of process microbes on soluble substrate - _{max (Ac)} h^–1 maximum specific growth rate constant for growth of process microbes on acetate - Ks _Ss kg/m³ saturation coefficient for growth of process microbes on soluble substrate - Ks _Ac kg/m³ saturation coefficient for growth of process microbes on acetate - K _d h^–1 death/lysis rate constant for process microbes - K _i kg/m³ inhibition constant for growth of process microbes on acetate - K _L h^–1 lysis rate constant for whole yeast cells - K _h h^–1 hydrolysis rate constant for particulate biomass     

A modified unstructured mathemathical model for the penicillin G fed-batch fermentation

Summary In this paper, an updated unstructured mathematical model for the penicillin G fed-batch fermentation is proposed, in order to correct some physical and biochemical shortcomings in the model of Heijnen et al. (1979,Biotechnol. Bioeng.,21, 2175–2201) and the model of Bajpai and Reuß (1980,J. Chem. Tech. Biotechnol.,30, 332–344). Its main features are the consistency for all values of the variables, and the ability to adequately describe different metabolic conditions of the mould. The model presented here can be considered as the translation of the latest advances in the biochemical knowledge of the penicillin biosynthesis.Nomenclature t time (h) - S amount of substrate in broth (g) - X amount of cell mass in broth (g) - P amount of product in broth (g) - V fermentor volume (L) - F input substrate feed rate (L/hr) - C _s S/V substrate concentration in broth (g/L) - C _x X/V cell mass concentration in broth (g/L) - C _P P/V product concentration in broth (g/L) - s _F substrate concentration in feed stream (g/L) - E _m parameter related to the endogenous fraction of maintenance (g/L) - E _p parameter related to the endogenous fraction of production (g/L) - K _x Contois saturation constant for substrate limitation of biomass production (g/g DM) - K _s Monod saturation constant for substrate limitation of biomss production (g/L) - K _p saturation constant for substrate limitation of product formation (g/L) - K _i substrate inhibition constant for product formation (g/L) - m _s maintenance constant (g/g DM hr) - k _h penicillin hydrolysis or degradation constant (hr^–1) - Y _x/s cell mass on substrate yield (g DM/g) - Y _p/s product on substrate yield (g/g) - specific substrate consumption rate (g/g DM hr) - specific growth rate (hr^–1) - _substr specific substrate to biomass conversion rate (hr^–1) - _x maximum specific substrate to biomass conversion rate (hr^–1) - specific production rate (g/g DM hr) - _p specific production constant (g/g DM hr)     

Nonlinear adaptive optimization of biomass productivity in continuous bioreactors

A novel on-line adaptive optimization algorithm is developed and applied to continuous biological reactors. The algorithm makes use of a simple nonlinear estimation model that relates either the cell-mass productivity or the cell-mass concentration to the dilution rate. On-line estimation is used to recursively identify the parameters in the nonlinear process model and to periodically calculate and steer the bioreactor to the dilution rate that yields optimum cell-mass productivity. Thus, the algorithm does not require an accurate process model, locates the optimum dilution rate online, and maintains the bioreactors at this optimum condition at all times. The features of the proposed new algorithm are compared with those of other adaptive optimization techniques presented in the literature [1–5]. A detailed simulation study using three different microbial system models [3, 6–7] was conducted to illustrate the performance of the optimization algorithm.List of Symbols A(q ^–1) polynomial in q ^–1 - b bias term - c _F nutrient cost term - B(q ^–1) polynomial in q ^–1 - C(q ^–1) polynomial in q ^–1 - CMPR kg/(m³ · h) cell mass productivity - D 1/h dilution rate - D _opt 1/h optimum dilution rate - E(q ^–1) polynomial in q ^–1 - h exponential filter constant - J objective function - k time index - K _m Monod's constant - n optimization interval - P covariance matrix - q ^–1 backward shift operator - r defined by equation (28) - S kg/m³ substrate concentration - S _F kg/m³ feed substrate concentration - T _s h sampling period - u vector containing previous input values - V dm³ fermenter volume - X kg/dm³ cell mass concentration - Y output variable - Y vector containing previous output values - Y _x/s g/g yield coefficient - optimization tuning constant - vector linear or nonlinear combination of u and Y - denominator covariance matrix update equation - forgetting factor - parameter vector - 1/h specific growth rate - _m 1/h maximum specific grow rate     

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)     

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     

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      

High yeast concentration in continuous fermentation with cell recycle obtained by tangential microfiltration

Summary Continuous fermentation fed by 150 kg/m³ of glucose with total cell recycling by tangential microfiltration enabled yeasts concentration of 300 kg/m³ (dry weight) to be reached with a dilution rate of 0,5h^–1 and a cell viability greater than 75%. The stability of this system was tested for 50 residence times of the permeate. The method can be used both for the production of cell concentrates and for high rates of metabolite production.Nomenclature D. W. dry weight - X_T (kg/m³) total cell concentration D.W. - X_V (kg/m³) viable cell concentration D.W. - V viability of cell culture in per cent of total cell concentration - S (kg/m³) glucose concentration - P (kg/m³) ethanol concentration - D (h) dilution rate - R (kg/kg) fermentation yield - (h) specific growth rate - vp(kg/kg/h) specific alcohol production rate - (m) yeast size - (kg/kg) kg of intracellular water per kg of dry cells     

Xylitol formation by Candida boidinii in oxygen limited chemostat culture

Summary Production of xylitol by Candida boidinii NRRL Y-17213 occurs under conditions of an oxygen limitation. The extent to which substrate is converted to xylitol and its coproducts (ethanol, other polyols, acetic acid), and the relative flow rates of substrate to energetic and biosynthetic pathways is controlled by the degree of oxygen limitation.With decrease in oxygen concentration in the inlet gas, for a constant dilution rate of 0.05 1/h. the specific oxygen uptake rate decreased from 1.30 to 0.36 mmol/gh Xylitol was not produced at specific oxygen uptake rates above 0.91 mmol/gh. Upon shift to lower oxygen rates, specific xylitol production rate increased more rapidly than specific ethanol production rate:Nomenclature D dilution rate (1/h) - DOT dissolved oxygen tension (%) - m_o2 maintenance coefficient (mmol O₂/g cell mass h) - q_o2 specific oxygen uptake rate (mmol O₂/g cell mass h) - q_s specific xylose uptake rate (g xylose/g cell mass h) or (mmol xylose/g cell mass h) - q_x specific xylitol production rate (g xylitol/ g cell mass h) or (mmol xylitol/ g cell mass h) - q_e specific ethanol production rate (g ethanol/ g cell mass h) or (mmol ethanol/ g cell mass h) - q_CO2 specific carbon dioxide production rate (mmol CO₂/g cell mass h) - S xylose concentration (g/1) - Y_cm/s cell mass yield coefficient, (g cell mass/mmol xylose) or (g cell mass/ g xylose consumed) - Y_cm/O2 cell mass yield coefficient, (g cell mass/mmol O₂) - Y_X/S xylitol yield coefficient (g xylitol/g xylose consumed) - Y_x/O2 xylitol yield coefficient (g xylitol/mmol O₂) - Y_e/s ethanol yield coefficient (g ethanol/g xylose consumed) - OUR oxygen uptake rate (mmol O₂/1h) - specific growth rate (1/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      

The scale-up of plant cell culture: Engineering considerations

The enormous versatility of plants has continued to provide the impetus for the development of plant tissue culture as a commercial production strategy for secondary metabolites. Unfortunately problems with slow growth rates and low products yields, which are generally non-growth associated and intracellular, have made plant cell culture-based processes, with a few exceptions, economically unrealistic. Recent developments in reactor design and control, elicitor technology, molecular biology, and consumer demand for natural products, are fuelling a renaissance in plant cell culture as a production strategy. In this review we address the engineering consequences of the unique characteristics of plant cells on the scale-up of plant cell culture.Abbreviations a gas-liquid interfacial area per volume - C dissolved oxygen concentration - C^* liquid phase oxygen concentration in equilibrium with the partial pressure of oxygen in the bulk gas phase - K_L overall mass transfer coefficient - k_L liquid film mass transfer coefficient - m_O2 cell maintenance coefficient for oxygen - OTR oxygen transfer rate - OUR oxygen uptake rate - pO₂ partial pressure of oxygen - STR stirred-tank reactor - v.v.m. volume of gas fed per unit operating volume of reactor per minute - X biomass concentration - Y_x/O2 biomass yield coefficient for oxygen - specific growth rate     

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     

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)     

Interrelation between penicillin productivity and growth rate

Summary Fermentations were carried out in an 801 tower-loop reactor with pellets of Penicillium chrysogenum. The development of the inner structure of the pellets with regard to various fermentation conditions was observed by means of histological preparations of the pellets. Under conditions of energy-source-limitation mycelial tip growth and lysis of other mycelial parts exist simultaneously. Thus the net growth rate (formation rate of cell mass) is higher than the gross growth rate (multiplication rate of cell mass). Under conditions of nitrogen limitation, gross growth rate and net growth rate are identical. A very strict correlation between gross growth rate and penicillin production rate was found as long as sufficient oxygen supply could be maintained and carbon catabolite repression was avoided. The energy source requirement of the biomass can be described with the sum of three terms that correspond to gross growth, lysis compensation growth and maintenance.Symbols a Constant 1/l h - b Constant - K Decay rate constant for product 1/h - K ₁ Substrate inhibition constant g/l - K _op Controls saturation constant for oxygen g/l - K _p Saturation constant for substrate g/l - m Maintenance coefficient 1/h - ms Apparent maintenance coefficient 1/h - O Dissolved oxygen concentration g/l - P Product concentration g/l - p Exponent of O - q Specific productivity 1/h - S Substrate concentration g/l - t Time h - t ₁ Beginning of production phase h - t ₂ Time of pellet dissolution h - V Liquid volume of fermentation broth l - X Dry cell mass concentration g/l - Y Yield of dry cell mass from energy substrate - g Specific gross growth rate of biomass 1/h - l Specific lysis rate of cell mass 1/h - n Specific net growth rate of cell mass 1/h - p Maximum specific rate of product formation 1/h