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.     

Mass transfer and liquid mixing in an airlift reactor with a net draft tube

Mass transfer and liquid mixing in an airlift reactor with a net draft tube were experimentally investigated. Four different column diameters were considered. The mass transfer was measured using the volumetric gas-liquid mass transfer coefficient which was determined by the dynamic method. The mass transfer coefficients in the airlift reactors with different column diameters were not always higher than those in the bubble columns. The liquid mixing was measured using mixing time which was determined by a pulse technique. Under the same superficial gas velocity, the mixing times of the airlift reactors with a net draft tube were always less than those of the bubble columns.List of Symbols C mol·dm^–3 bulk concentration of dissolved oxygen - C ₀ mol·dm^–3 initial concentration of dissolved oxygen - C _e mol·dm^–3 saturated concentration of dissolved oxygen - ¯C dimensionless dissolved oxygen concentration - D _c cm diameter of column - D _N cm diameter of the nozzle hole - D _T cm diameter of the net draft tube - H _L cm static liquid height - H _T cm height of the net draft tube - k _L a hr^–1 volumetric mass transfer coefficient - L _T cm length of the net draft tube - t _M sec mixing time of the liquid phase - t ₀ sec mixing time of the liquid phase in a bubble column - V _L dm³ volume of the liquid phase - U _g cm/s superficial air velocity     

Physical characterisations of a single-stage Kühni-type aqueous two-phase extraction column

The main parameters which influence the behaviour of phase separation in a single-stage Kühni-type aqueous two-phase extraction column containing polyethylene (PEG) and di-potassium hydrogen phosphate were characterised. Two aqueous two-phase system (ATPS) composed of 12% (w/w) PEG 1450 and 12% (w/w) di-potassium hydrogen phosphate (designated as 12/12) and 12% (w/w) PEG 1450 and 11% (w/w) di-potassium hydrogen phosphate (designated as 12/11) were chosen in this study. The hold-up ɛ_D increased with increasing impeller speeds and mobile phase flow rates. Phase separation for the 12/11 system was slower than that for the 12/12 system, which resulted in higher dispersed phase hold-up values for the 12/11 system. For 12/12 system, mass transfer of plasmid DNA (pDNA) from the dispersed mobile phase to the stationary phase increased rapidly with increasing impeller speeds of 130, 160 and 200 rpm which was reflected in the decreased values for C_T/C_To. The degree of back-mixing quantified by the axial dispersion coefficient D_ax was estimated to be 2.7 × 10⁻⁶ m² s⁻¹.     

Emission of microorganisms from biofilters

Experiments are reported on the discharge of microbial germs by biofilter systems used for the treatment of waste gases containing volatile organic compounds. The systems investigated concern six full-scale filter installations located in the Netherlands in several branches of industry, as well as a laboratory-scale installation used for modelling the discharge process. It is concluded that the number of microbial germs (mainly bacteria and to a much smaller extent moulds) in the outlet gas of the different full scale biofilters varies between 10³ and 10⁴ m^–3, a number which is only slightly higher than the number encountered in open air and of the same order of magnitude encountered in indoor air. It is furthermore concluded that the concentration of microorganisms of a highly contaminated inlet gas is considerably reduced by the filtration process. On the basis of the experiments performed in the laboratory-scale filter bed, it is shown that the effect of the gas velocity on the discharge process results from two distinctive mechanisms: capture and emission. A theoretical model is presented describing the rate processes of both mechanisms. The model presented and the experimentally determined data agree rather well.List of Symbols a _s m^–1 specific area of the packing material - C m^–3 microbial gas phase concentration - C _e , C _i m^–3 microbial concentration in the exit and inlet gas resp. - CFU colony-forming-units - d _c , d _m m diameter of collecting and captured particle resp. - D m diameter of the filter bed - E single particle target efficiency - H m bed height - k _c s^–1 first order capture rate constant per unit of bedvolume - k _e m^–3 emission rate constant per unit of bedvolume - n number of observations - r _c , r _e m^–3 s^–1 capture and emission rate per unit of bed-volume - Re = Reynolds number - S _t = Stokes number - u m s^–1 superficial gas velocity - u _m m s^–1 superficial gas velocity at which C _e = C _i Greek Symbols void fraction of the filter bed - kg m^–3 density of the gas phase - _m kg m^–3 density of captured particle - Pa s dynamic gas phase viscosity - = filter bed efficiency     

Effect of serum concentration on production of monoclonal antibodies and on shear sensitivity of a hybridoma

The effect of serum content of culture medium on the specific production rate of monoclonal antibodies (Mab's) and on shear sensitivity has been studied with hybridoma's, cultured in a continuous stirred tank reactor (CSTR). No decrease in specific Mab-production was found when the serum concentration was reduced from 10 to 2.5%, while steady state cell concentrations were hardly affected as well. In contrast the cell death rate in a bubble column strongly increased when the serum concentration was lowered, which could be ascribed to a reduced physical protective effect by the serum.List of Symbols k _d s^–1 death-rate constant - k _g s^–1 growth-rate constant - D m diameter of bubble column - d m diameter of air bubble - H m height of bubble column - X m³ hypothetical killing volume - X (m³/m³) specific hypothetical killing volume - F m³/s volumetric air-flow rate - C cells/m³ number of viable cells - C₀ cells/m³ number of viable cells at t=0 - t s time     

Cell recovery by continuous flotation

Summary Cell recovery by means of continuous flotation of the Hansenula polymorpha cultivation medium without additives was investigated as a function of the cultivation conditions as well as of the flotation equipment construction and flotation operational parameters. The cell enrichment and separation is improved at high liquid residence times, high aeration rates, small bubble sizes, increasing height of the aerated column, and diameter of the foam column. Increasing cell age and cultivation with nitrogen limitation reduce the cell separation.Symbols C_P cell mass concentration in medium g·l^–1 - C_R cell mass concentration in residue g·l^–1 - C_S cell mass concentration in foam liquid g·l^–1 - V equilibrium foam volume cm³ - V gas flow rate through the aerated liquid column cm³·s^–1 - V_F feed rate to the flotation column ml/min - ¹ V _S/V foaminess s - mean liquid residence time in the column s     

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     

A dynamic cell cycling model for growth of baker's yeast and its application in profit optimization

A cell cycling model for unequal budding yeast Saccharomyces cerevisiae is proposed and verified by steady state data from experiments available in the literature. This model can be used to determine the relative fraction of the cells in any cycling phase or with any genealogical age during fermentation. As the quality of yeast is strongly influenced by the cycling process, the model could therefore be used to control the quality of the harvested yeast cells. The input of the cell cycling model is the specific growth rate , which is obtained from a metabolic model for S. cerevisiae proposed earlier. With this extended model system not only the quality control, but also the whole economical profit optimization can be carried out. Simulations were done to optimize the profit of a commercial scale baker's yeast production process by manipulating substrate feeding rate and substrate concentration under different aeration rates, fermentation periods and other conditions applied in industry.List of Symbols B h budding phase - C _d1, C _d2, C _p1' parameters in cycling phase equations - C _p2, C _b1, C _b2, d _s m Sauter-diameter - E kg/m³ ethanol concentration - E1, E2 state variables in the metabolic model - E _G mean relative gas hold-up - f parameter vector of the regulation model - F system matrix of the regulation model - F or F(t) m³/h substrate feeding rate - Fr Froude number - FBC, FDC, FPC % fraction of daughter cells, unbudded daughter cells and unbudded parent cells - g m/s² acceleration of gravity - K _B1–3, K _EG parameters in metabolic model - K ₃, K _Ad , L ₃ K ₃ E, K_O, K_S limitation constants for ethanol, oxygen and substrate - k _La h^–1 volumetric mas transfer coefficient - m _ATP mol(gh)^–1 maintenance coefficient - nb, nd, np number of cycling age intervals in budding cycling phase, unbudded daughter cycling phase and unbudded parent cycling phase - Nt number of total cells - OF mg/dm³ concentration of dissolved oxygen - P kg total yeast product in dry weight - P/O effectiveness of oxidative phosphorylation - q _O20 mol(gh)^–1 minimum specific oxygen uptake ability - q _O2 mol(gh)^–1 specific oxygen uptake rate - q _O2max mol(gh)^–1 maximum q _O2 given by metabolic regulation - q _s mol(gh)^–1 specific glucose uptake rate - q _Smax mol(gh)^–1 maximum q _S - R(·) switch function - r _Ac mol(gh)^–1 specific acetyl-CoA-consumption rate - r _Acmax saturation rate of r _Ac - r _E1 mol(gh)^–1 specific ethanol production rate - r _E2 mol(gh)^–1 specific ethanol uptake rate - r _SO mol(gh)^–1 minimum value of r _Smax - r _s mol(gh)^–1 specific rate of glycolysis - r _Smax mol(gh)^–1 maximum specific rate of gluconeogenesis given by metabolic regulation - S kg/m³ total reduced sugar concentration - S _R kg/m³ substrate concentration in feed - T h cell number doubling time - T _f h fermentation period - Ud h unbudded daughter phase - Up h unbudded parent phase - V _F m³ volume of liquid phase in fermentor - V _G m³/h aeration rate - w _sg m/s superficial gas velocity - X kg/m³ dried cell concentration - Y _ATP g(molATP)^–1 yield coefficient of ATP - z state vector in regulation model - the factor of fermentative activity decrease caused by budding cells - or(t) h^–1 specific growth rate - h discrete unit of cycling age     

Optimisation of agitation and aeration in fermenters

The problem of optimising agitation and aeration in a given fermenter is addressed. The objective function is total electric power consumed for agitation, compression and refrigeration. The major constraint considered is to ensure that the dissolved oxygen concentration is above the critical value. It is shown that it is possible to analytically calculate the optimal pair (air flowrate, stirrer speed) and that, at least for the industrial antibiotics fermentation used as case-study, the optimum lies within a window for satisfactory operation, limited by other possible constraints to the problem. Savings achievable by optimal operation as compared with current industrial procedure were found to be around 10% at pilot plant scale (0.26 m³) and 20% at full scale (85 m³).List of Symbols A fermenter cross sectional area (m²) - C dissolved oxygen concentration (mole m^–3) - C ^* DO concentration in equilibrium with the gas (mole m^–3) - C _crit critical DO concentration (mole m^–3) - C _p specific heat of air at constant pressure (J kg^–1 K^–1) - C _sp dissolved oxygen set point (mole m^–3) - C _v specific heat of air at constant volume (J kg^–1 K^–1) - D agitator diameter (m) - f pressure correction of air flow-rate - (Fl _g)_F aeration number at flooding - (Fr _g)_F froude number at flooding - k coefficient in expression for mass transfer coefficient - K _La volumetric oxygen transfer coefficient (s^–1) - m power exponent in expression for mass transfer coefficient - n gas flow rate exponent in expression for mass transfer coefficient - n ^* number of impellers - N rotation speed (s^–1) - N _F rotation speed at flooding (s^–1) - N _p unaerated power number - N _pg aerated power number - OUR Oxygen Uptake Rate (mole m^–3 s^–1) - p ₀ atmospheric pressure (N m^–2) - p ₁ compressor exit pressure (N m^–2) - p ₂ pressure at the bottom of the fermenter (N m^–2) - p ₃ pressure at the top of the fermenter (N m^–2) - P _c compression power (W) - P _d power added by expansion (W) - P _ev power removed by evaporation (W) - P _g agitation power (W) - P _m power added by metabolism (W) - P _r power removed by refrigeration (W) - P _t total power (W) - Q air flow-rate at atmospheric conditions (m³ s^–1) - Q _f air flow-rate at average fermenter conditions (m³ s^–1) - s ₀ absolute humidity at atmospheric conditions - s ₃ absolute humidity at fermenter exit - T tank diameter (m) - V liquid volume (m³) - v _s gas superficial velocity (m s^–1) - _i parameter defined in the text - safety margin for dissolved oxygen (mole m^–3) - ratio of specific heats of air - _g agitation efficiency - _c compression efficiency - _r refrigeration efficiency - liquid density (kg m^–3) - _g air density (kg m^–3) - latent heat of vaporisation of water (J kg^–1) The authors are grateful to Elsa Silva, Carlos Lopes, Carlos Aguiar, Fernando Mendes, and Alexandre Cardoso, who helped with parts of this work, and to CIPAN for permission to publish these data.     

Hydrodynamics and mass transfer in an airlift loop fermentor

Summary The hydrodynamics and mass transfer behaviour of an airlift fermentor with an external loop (height 10m) has been investigated by measuring gas and liquid velocities, gas hold-up, liquid mixing and oxygen transfer coefficients. Liquid phase properties, i.e., ionic strength, viscosity and surface tension have been varied by altering the fermentation media. Results are compared with those from bubble column experiments performed in the same unit. It is shown, that more uniform two-phase flow in the airlift leads to advantages in scale-up and operation.Nomenclature a Specific interfacial area per volume of dispersion (m²/m³) - c Local concentration of tracer (kmol/m³) - c Concentration of tracer at infinite time (kmol/m³) - C_L Concentration of oxygen in the liquid bulk (kmol/m³) - C_L ^* Concentration of oxygen in the interface (kmol/m³) - D_ax Axial dispersion coefficient (cm²/s) - I Ionic strength (kmol/m³) - i Inhomogeneity [defined in Eq. (2)] - Rate of oxygen transfer (kmol/s) - t_c Circulation time (s) - t_M Mixing time (s) - V_R Volume of gas-liquid dispersion (m³) - V_SG Superficial gas velocity in up-flow column (m/s) Greek letter symbols _L Oxygen transfer coefficient (m/s) - Dynamic viscosity (m Pa s) - Surface tension (m N/m) Presented at the First European Congress on Biotechnology, Interlaken, September 25–29, 1978     

Biological elimination of volatile xenobiotic compounds in biofilters

Biofiltration is a technique which is frequently applied for the odour abatement of waste gases. This technique is based on the ability of microorganisms (generally bacteria, and to a small extent moulds and yeasts) to degrade several organic as well as inorganic compounds to mineral end-products, like water and carbon dioxide. In the case of biofiltration, microorganisms are attached to suited packing materials in the filter, which contain the inorganic nutrients necessary for microbial growth. In order to make biofiltration applicable on a larger scale in process industry, it is necessary to find microorganisms able to eliminate compounds which are strange to life, the so-called xenobiotics. At the Eindhoven University of Technology microorganisms were isolated for the elimination of a number of xenobiotics, e.g. aromatic compounds and chlorinated hydrocarbons.Both the microkinetics of the biodegradation in aqueous batch systems and the macrokinetics of biofilters were studied. Special attention was paid to the influence of the superficial gas velocity and the organic load on the filter bed's elimination capacity. Also the discontinuous operation of biofilters is discussed.Nomenclature C ₁],C _1,0,C _í mol/m³ liquid phase concentration - K _s mol/m³ saturation constant - Y kg/kmol yield coefficient - t, t d time - _md^–1 maximum growth rate - X, X ₀ g/m³ microorganism concentration - P _gPa partial vapour pressure - H Pa m³/mol Henry's law constant - m (-) distribution coefficient - C _g,C _g,max mol/m³ gas phase concentration - R J (mol K) gas constant - T K temperature - (-) gas phase- and liquid phase volume ratio - m³/(m² h) superficial gas velocity - H m height of the filter bed - h m height in the filter bed In memory of Dick van Zuidam     

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     

Oxygen transfer in an airlift reactor with multiple net draft tubes

A pilot scale airlift reactor with multiple net draft tubes was developed to improve oxygen transfer in the reactor. The reactor was 0.29 m in diameter and 2 m height. A steadystate sulfite oxidation method was applied to determine an overall volumetric mass transfer coefficient. Oxygen transfer of the proposed airlift reactor can be 60–100% higher than that of bubble columns under the same operating conditions.List of Symbols C ^* mol·dm^–3 saturated concentration of dissolved oxygen - C _L mol·dm^–3 bulk concentration of dissolved oxygen - G mol/min nitrogen flow rate - k _L a hr^–1 the volumetric gas-liquid mass transfer coefficient - Mo ₂ g/mol molecular weight of oxygen - OTR g/min the oxygen transfer rate - U _g cm/s superficial air velocity - V _L dm³ volume of the liquid phase - _in oxygen mole ratio in the inlet gas - _out oxygen mole ratio in the outlet gas     

Removal of toluene in waste gases using a biological trickling filter

The removal of toluene from waste gas was studied in a trickling biofilter. A high level of water recirculation (4.7 m h^–1) was maintained in order to keep the liquid phase concentration constant and to achieve a high degree of wetting. For loads in the range from 6 to 150 g m^–3 h^–1 the maximum volumetric removal rate (elimination capacity) was 35±10 g m^–3 h^–1, corresponding to a zero order removal rate of 0.11±0.03 g m^–2 h^–1 per unit of nominal surface area. The surface removal was zero order above the liquid phase concentrations of approximately 1.0 g m^–3, corresponding to inlet gas concentrations above 0.7–0.8 g m^–3. Below this concentration the surface removal was roughly of first order. The magnitude of the first order surface removal rate constant, k_1A , was estimated to be 0.08–0.27 m h^–1 (k_1A a=24–86 h^–1). Near-equilibrium conditions existed in the gas effluent, so mass transfer from gas to liquid was obviously relatively fast compared to the biological degradation. An analytical model based on a constant liquid phase concentration through the trickling filter column predicts the effluent gas concentration and the liquid phase concentration for a first and a zero order surface removal. The experimental results were in reasonable agreement with a very simple model valid for conditions with an overall removal governed by the biological degradation and independent of the gas/liquid mass transfer. The overall liquid mass transfer coefficient, K_La, was found to be a factor 6 higher in the system with biofilm compared to the system without. The difference may be explained by: 1. Difference in the wetting of the packing material, 2. Mass transfer occurring directly from the gas phase to the biofilm, and 3. Enlarged contact area between the gas phase and the biofilm due to a rough biofilm surface.     

Cross-flow filtration as a method of separating fungal cells and purifying the polysaccharide produced

Cross-flow filtration (CFF) has been investigated as a method of separating filamentously growing fungal cells and purifying the polysaccharide produced. The effects of transmembrane pressure, module geometry (e.g. channel height or tube diameter), tangential feed velocity and cell as well as polysaccharide concentration are discussed. Apart from these experiments, influences by the recirculation pump used are shown.List of Symbols b _f fouling index - b factor refering to the behaviour of the sublayer - C kg · m^–3 concentration - C _g kg · m^–3 solute concentration at the membrane - C _b kg · m^–3 solute concentration in the bulk phase - D s_-1 shear rate - k m · s^–1 mass-transfer coefficient - K mPa · sⁿ consistency index - n flow behaviour index - P _w m³ · s^–1 · m^–2 rate of permeation - P _w1 m³ · s^–1 · m^–2 rate of permeation at 1 minute - P _w m³ · s^–1 · m^–2 rate of permeation at the beginning - p Pa pressure - Q m² largest cross-section of a particle - q m² smallest cross-section of a particle - Re Reynolds number - R _f ^–1 fouling resistance - R _m m^–1 membrane resistance - t s time - w m · s^–1 tangential feed velocity Greek Symbols friction factor - pTM Pa transmembrane pressure - mPa · s shear viscosity - _sp specific viscosity (rel. increase of viscosity _sp=_rel-1) - [] m³· kg^–1 intrinsic viscosity - _w m² · s^–1 kinematic viscosity - kg · m^–3 density Indices b bulk - cell cells - f fouling - g gelling - PS polysaccharide - rel relative - sp specific - w water     

Analysis of continuous fermentation processes in aqueous two-phase systems

Simulations of continuous ethanol or acetonobutylic fermentations in aqueous two-phase systems show that at high substrate feed concentrations it is possible to obtain solvent productivities about 25–40% higher than in conventional systems with cell recycle if the biomass bleed rate is kept about one tenth of the value of D.List of Symbols a Volumetric fraction of dextran rich phase - B h^–1 Bleed rate - D h^–1 Dilution rate - P kg m^–3 Product concentration - PD kg m^–3 h^–1 Productivity - S kg m^–3 Substrate - X kg m^–3 Biomass - Partition coefficient     

Apparent viscosity for non-Newtonian fermentation media in bioreactors

The apparent viscosity of non-Newtonian fermentation media is examined. The present state of this subject is discussed. The energy dissipation rate concept is used for a new evaluation of the apparent viscosity in bioreactors, i.e. stirred tank and bubble column bioreactors. The proposed definition of the apparent viscosity is compared with the definitions available in the literature.List of Symbols A _d m ² downcomer cross-sectional area - A _r m ² riser cross-sectional area - a m^–1 specific surface area - C constant in eq. (13) - D m column diameter - D _I m impeller diameter - g m s^–2 gravitational acceleration - h J m^–2 s^–1 K^–1 heat transfer coefficient - K Pa s ⁿ consistency index in a power-law model - k constant in eq. (3) - k _L m s ^–1 liquid-phase mass transfer coefficient - N s^–1 impeller speed - n flow index in a power-law model - P W power input - Re Reynolds number ND _I ^/2 /(/) - U _sg m s ^–1 superficial gas velocity - (U _sg ) _r m s^–1 superficial gas velocity based on riser - V-m³ liquid volume - v ₀ m s^–1 friction velocity Greek Symbols s^–1 shear rate - _c s^–1 characteristic shear rate - W kg^–1 energy dissipation rate per unit mass - W kg^–1 characteristic energy dissipation rate per unit mass - Pa s viscosity - _app Pa s apparent viscosity - kg m^–3 density - Pa shear stress     

Effects of downcomer-to-riser cross sectional area ratio on operation behaviour of external-loop airlift bioreactors

Experiments performed in two external-loop airlift bioreactors of laboratory and pilot scale, (1.880–1.189) · 10^–3 m³ and (0.170-0.157)m³, respectively, are reported. The A _D /A _R ratio was varied between 0.111–1.000 and 0.040–0.1225 in the laboratory and pilot contractor respectively.Water and solutions of different coalescence (2-propanol 2% vol, 1 M Na (glucose 50% wt/vol) and rheological behaviour (non-Newtonian starch solutions with consistency index K=0.061–3.518 Pas ⁿ and flow behaviour index n=0.86-0.39), respectively, were used as liquid phase. Compressed air at superficial velocities v _SGR =0.016–0.178 ms^–1 in the laboratory contactor and v _SGR =0.010–0.120 ms^–1 in the pilot contactor, respectively was used as gaseous phase.The A _D /A _R ratio affect gas-holdup behaviour as a result of the influence of A _D /A _R on liquid circulation velocity.Experimental results show that A _D /A _R ratio affect circulation liquid velocity by modifying he resistence to flow and by varying the fraction of the total volume contained in downcomer and riser. A _D /A _R ratio has proven to be the main factor which determines the friction in the reactor. Mixing time increases with increasing of the reactor size and decreases with A _D /A _R decreasing.The volumetric gas-liquid mass transfer coefficient increases with A _D /A _R ratio decreasing, as a result of variations of the liquid velocity with A _D /A _R , which affect interfacial areas.Correlations applicable to the investigated contactors have been presented, together with the fit of some experimental data to existing correlation in literature.List of Symbols A _D downcomer cross sectional area (m²) - A _R riser cross sectional area (m²) - a coefficient in Eq. (9) (-) - a _L gas-liquid interfacial area per unit volume (m^–1) - b coefficient in Eq. (9) (-) - C tracer concentration (kg m^–3) - C tracer concentration at the state of complete mixing (kg m^–3) - c coefficient in Eq. (12) - c _S coefficient in Eq. (5) - D _D downcomer diameter (m) - D _R riser diameter (m) - d _B bubble size (m) - H _D downcomer height (m) - H _d dispersion height (m) - H _L gas-free liquid height (m) - H _R riser height (m) - I inhomogeneity (-) - K consistency index (Pa s ⁿ ) - k _L a volumetric gas-liquid oxygen mass transfer coefficient (s^–1) - m exponent in Eq. (12) (-) - n flow behaviour index (-) - P _G power input due to gassing (W) - t _M mixing time (s) - V _A connecting pipe volume (m³) - V _D downcomer volume (m³) - V _d volume of dispersion (m³) - V _R riser volume (m³) - V _T total reactor liquid volume (m³) - v _SGR riser gas superficial velocity (m s^–1) - _GR riser gas holdup (-) - shear rate (m s^–1) - _app apparent viscosity (Pa s) - shear stress     

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      

Fully automated high-performance liquid chromatography of ciprofloxacin with direct injection of plasma and Mueller–Hinton broth for pharmacokinetic/pharmacodynamic studies

An isocratic high-performance liquid chromatographic method with column switching and direct injection has been developed to determine ciprofloxacin in plasma and Mueller–Hinton broth. An on-line dilution of the sample was performed with a loading mobile phase consisting of 173 mM phosphoric acid. The analyte was retained on a LiChrocart 4-4 precolumn filled with a LiChrospher 100 RP18, 5 μm. An electric-actuated system with two six-port valves allowed a clean-up step with a mixture 20 mM phosphate buffer (pH 3.5)–methanol (97: 3, v/v) and the transfer of the analyte by a back-flush mode to a 150×4.6 mm I.D. column packed with a Kromasil C₈ 5 μm, using a mobile phase of 20 mM phosphate buffer (pH 3.5)–acetonitrile (85:15, v/v). Fluorescence detection allowed a quantification limit of 0.078 μg/ml with a 40-μl sample size. The method was evaluated to determine its usefulness in studying the pharmacokinetic/pharmacodynamic behaviour of ciprofloxacin in an in vitro model.