Concentric-tube airlift bioreactors

Gas holdup investigations were performed in three concentric-tube airlift reactors of different scales of operation (RIMP: 0.070 m³; RIS-1: 2.5 m³; RIS-2: 5.2 m³; nominal volumes). The influences of the top and bottom clearances and the flow resistances at the downcomer entrance were studied using tap water as liquid phase and air as gaseous phase, at atmospheric pressure. It was found that the gas holdup in the individual zone of the reactor: riser, downcomer and gas-separator, as well as that in the overall reactor is affected by the analyzed geometrical parameters in different ways, depending on their effects on liquid circulation velocity. Gas holdup was satisfactorily correlated with Fr, Ga, bottom spatial ratio (B), top spatial ratio (T), gas separation ratio (Y) and downcomer flow resistance ratio (A _d /A _R ). Correlations are presented for gas holdup in riser, downcomer, gas separator and for the total gas holdup in the reactor. All the above stressed the importance of the geometry in dynamic behaviour of airlift reactors.     

Modelling mixing parameters in concentric-tube airlift bioreactors

Axial dispersion of the liquid phase was investigated in a concentric-tube airlift bioreactor (RIMP: V _L=0.70?m³) as a whole and in the separate zones (riser, downcomer, gas-separator) using the axial dispersion model. The axial dispersion number Bo and the axial dispersion coefficient, D _ax were determined from the output curves to an initial Dirac pulse, using the tracer response technique. They were analyzed in relation to process and geometrical parameters, such as: gas superficial velocity, ν_SGR; top clearance, h _S; bottom clearance, h _B, and resistances at downcomer entrance expressed as A _d/A _R ratio. Correlations between Bodenstein numbers in the overall bioreactor and riser and downcomer sections (Bo_T,Bo_R,Bo_D) and the geometrical and process parameters were developed, which can allow to assess the complex influence of these parameters on liquid axial dispersion.     

Modelling mixing parameters in concentric-tube airlift bioreactors

The mixing behaviour of the liquid phase in concentric-tube airlift bioreactors of different scale (RIMP: V_L=0.070 m³; RIS-1: V_L=2.50 m³; RIS-2: V_L=5.20 m³) in terms of mixing time was investigated. This mixing parameter was determined from the output curves to an initial Dirac pulse, using the classical tracer response technique, and analyzed in relation to process and geometrical parameters, such as: gas superficial velocity, x_SGR; top clearance, h_S; bottom clearance, h_B, and ratio of the resistances at downcomer entrance, A_d/A_R. A correlation between the mixing time and the specified operating and geometrical parameters was developed, which was particularized for two flow regimes: bubbly and transition (x_SGRА.08 m/s) and churn turbulent flow (x_SGR> 0.08 m/s) respectively. The correlation was applied in bioreactors of different scale with a maximum error of ᆲ%.     

Study of the liquid circulation velocity in external-loop airlift bioreactors

Liquid circulation velocity was studied in externalloop air-lift bioreactors of laboratory and pilot scale, respectively for different gas input rates, downcomer-to-riser cross-sectional area ratio, A _D/A_R and liquid phase apparent viscosities.It was found that, up to a gas superficial velocity in the riser v _SGR 0.04 m/s the dependency of v _SLR on v _SGR is in the following form: v _SLR = a v _SGR ^b , with the exponent b being 0.40. Over this value of v _SGR, only a small increase in liquid superficial velocity, v _SLR is produced by an increase in v _SGR. A _D/A_R ratio affects the liquid superficial velocity due to the resistance in flow and overall friction.For non-Newtonian viscous liquids, the circulation liquid velocity in the riser section of the pilot external-loop airlift bioreactor is shown to be dependent mainly on the downcomer-to-riser cross-sectional area ratio, A _D/A_R, the effective (apparent) liquid viscosity, _eff and the superficial gas velocity, v _SGR.The equation proposed by Popovic and Robinson [11] was fitted well, with an error of ± 20%.List of Symbols A _D m² downcomer cross-sectional area - A _Rm² riser cross-sectional area - a = coefficient in Eq. (7) - b = exponent in Eq. (7) - c _s m^–1 Coefficient in Eq. (3) - D _D m downcomer diameter - D _R m riser diameter - g m²/s gravitational acceleration - H _D m dispersion height - H _L m ungassed liquid height - K Pa s ⁿ consistency index - K _B = friction factor at the bioreactor bottom - K _F = friction factor - K _T = friction factor at the bioreactor top - V _L m³ liquid volume in the bioreactor - V _D m³ liquid volume in downcomer - V _R m³ liquid volume in riser - v _LDm/s downcomer linear liquid velocity - v _LR m/s riser linear liquid velocity - v _SGR m/s riser superficial liquid velocity - v _SLR m/s riser superficial liquid velocity - s^–1 shear rate - _GD = downcomer gas holdup - _GR = riser gas holdup - _eff Pa s effective (apparent) viscosity - Pa shear stress The authors wish to thank Mrs. Rodica Roman for the help in experimental data collection and to Dr. Stefanluca for the financial support.     

Concentric-tube airlift bioreactors

Mass transfer coefficients were measured in three concentric-tube airlift reactors of different scales (RIMP, V _L =0.07 m³;RIS?1,V _L =2.50 m³;RIS?2, V _L =5.20 m³). The effects of top and bottom clearance and flow resistances at downcorner entrance were studied in water-air system. Experimental results show that h _s ,h _B and A _d /A _R ratio affect K _L a values as a result of their influence on gas holdup and liquid velocity. The gas-liquid mass-transfer coefficients for all the geometric variables were successfully correlated as Sherwood number with Froude and Galilei numbers, the bottom spatial ratio (B=h _B /D _R ), the top spatial ratio , the gas separation ratio and the downcomer flow resistance ratio (R=A _d /A _R ). The proposed empirical model satisfactorily fitted the experimental data obtained in large airlift reactors and some data presented in literature.     

Residence time distribution of liquid phase in an external-loop airlift bioreactor

The residence time distribution analysis was used to investigated the flow behaviour in an external-loop airlift bioreactor regarded as a single unit and discriminating its different sections. The experimental results were fitted according to plug flow with superimposed axial dispersion and tank-in-series models, which have proved that it is reasonable to assume plug flow with axial dispersion in the overall reactor, in riser and downcomer sections, as well, while the gas separator should be considered as a perfectly mixed zone. Also, the whole reactor could be replaced with 105-30 zones with perfect mixing in series, while its separate zones, that is the riser with 104-27, the downcomer with 115-35 and the gas separator with 25-5 perfectly mixed zones in series, respectively, depending on gas superficial velocity, A_D/A_R ratio and the liquid feed rate.List of Symbols A _D cross sectional area of downcomer (m²) - A _R cross sectional area of riser (m²) - A ₁ A ₂ length of connecting pipes (m) - Bo Bodenstein number (Bo=vL·L/D _ax (-) - C concentration (kg m^–3) - C residence time distribution function - C ₀ coefficientEquation (12) - C _r dimensionless concentration - D _D diameter of downcomer (m) - D _R diameter of riser column (m) - D _ax axial dispersion coefficient (m²s^–1) - H _d height of gas-liquid dispersion (m) - H _L height of clear liquid (m) - i number of complete circulations - L length of path (m) - m order of moments - N _eq number of perfectly mixed zones in series - n _c circulating number - Q _c recirculating liquid flow rate (m³ s^–1) - q _F liquid feed flow rate (m³s^–1) - Q _G gas flow rate (m³s^–1) - Q _T total liquid flow rate (m³s^–1) - r recycle factor - s exponent inEquation (12) regarded as logarithmic decrement of the oscillating part of RTD curve - t time (s) - t _C circulation time (s) - t _s mean residence time (s) - t ₉₉ time necessary to remove 99% of the tracer concentration (s) - V _A volume of connecting pipes (m³) - V _D volume of downcomer (m³) - V _L liquid volume in reactor (m³) - V _R volume of riser (m³) - V _LD linear liquid velocity in downcomer (m s^–1) - V _LR linear liquid velocity in riser (m s^–1) - V _SLD superficial liquid velocity in downcomer (m s^–1) - V _SLR superficial liquid velocity in riser (m s^–1) - x independent variable inEquation (1) - ¯x mean value of x - z axial coordinate - _GR gas holdup in riser - _m(x) central moment of m order - ² variance - dimensionless time     

Hydrodynamics in external-loop airlift bioreactors with static mixers

Liquid circulation superficial velocity and gas holdup behaviours were investigated in an external-loop airlift bioreactor of 0.170?m³ liquid volume in gas-induced and forced-circulation-loop operation modes, in the presence of static mixers made of corrugated stainless steel pieces, resulting in packets with the height-to-diameter ratio equal to unity and using non-Newtonian starch solutions as liquid phase. The static mixers were disposed in the riser in three blocks, each with three mixing packets, successively turned 90° to the adjacent mixing element. It was found that in the presence of static mixers and forced-loop operation mode, liquid circulation superficial velocity in the riser section was significantly diminished, while gas holdup increased in a great measure. It was considered that static mixers split the fluid into individual streams and break up the bubbles, resulting in small bubble sizes with a relative homogeneous bubble distribution over riser cross section. They act as supplementary resistances in liquid flow, reducing riser cross sectional area, equivalent with A _D /A _R area ratio diminishing.     

Hydrodynamics of non-Newtonian liquids in external-loop airlift bioreactors

Gas holdup investigation was performed in two external-loop airlift bioreactors of laboratory (V _L =1.189·10^?3? 1.880·10^?3 m³; H _R =1.16 ? 1.56 m; H _D = 1.10 m; A _D /A _R = 0.111 ? 1.000) and pilot scale (V _L =0.157?0.170 m³; H _R =4.3?4.7 m; H _D =4.0?4.4 m;A _D /A _R =0.04?0.1225), respectively, using as liquid phase non-Newtonian starch solutions of different concentration with K=0.061?3.518 Pa sⁿ and n=0.86?0.39 and fermentation broths of P. chrysogenum, S. griseus, S. erythreus, B. licheniformis and C. acremonium at different hours since inoculation and from different batches. The influence of bioreactor geometry, liquid properties and the amount of introduced compressed air was investigated. The effect of sparger design on gas holdup was found to be negligible. It was found that gas holdup depends on the flow media index, ?_GR decreasing with the increase of liquid pseudoplasticity, A _D /A _R ratio and H _R /H _D ratio. The experimental data are in agreement with those presented in literature by Popovic and Robinson, which take into account liquid properties, geometric parameters and gas superficial velocity, with a maximum error of ±30%. It was obtained a correlation for gas holdup estimation taking into account the non-Newtonian behaviour of the fermentation broths and the dry weight of the solid phase, as well. The concordance between the experimental data and those calculated with the proposed correlation was good, with a maximum error of ±17%. Also, a dimensionless correlation for gas holdup involving superficial velocities of gas and liquid, cross sectional areas ratio, dispersion height to riser diameter ratio, as well as Froude and Morton numbers, was obtained.     

Hydrodynamics of non-Newtonian liquids in external-loop airlift bioreactors

In order to obtain further information on the behaviour and optimal design of external-circulation-loop airlift bioreactors, the liquid circulating velocity was studied using highly viscous pseudoplastic solutions of starch and antibiotic biosynthesis liquids of Penicillium chrysogenum, Streptomyces griseus, Streptomyces erythreus, Bacillus licheniformis and Cephalosporium acremonium. Measurements of liquid circulation velocity were made in laboratory and pilot plant external-loop airlift bioreactors, under various conditions concerning gas flow rate, riser liquid height at constant downcomer height, A _D /A _R ratio, using the impulse-response technique. It has been found that these parameters had a significant effect on liquid circulation velocity together with the apparent viscosity and dry weight of the solid phase in the biosynthesis liquids. For the tested liquids, the superficial liquid velocity in the riser section of an external-loop airlift bioreactor may be described by the following equation: where the exponents and the constant c take different values depending on the liquid phase properties and flow regime.     

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     

Saccharomyces cerevisiae fermentations in a pilot scale airlift bioreactor: Comparison of air sparger configurations

Hydrodynamic and oxygen transfer comparisons were made between two ring sparger locations, draft tube and annulus, in a concentric pilot scale airlift reactor with a baker's yeast suspension. Sectional hydrodynamic measurements were made and a mobile DOT probe was used to characterise the oxygen transfer performance through the individual sections of the reactor. The hydrodynamic performance of the reactor was improved by using a draft tube ring sparger rather than the annulus ring sparger. This was due to the influence of the ratio of the cross sectional area of the downcomer and riser (A _D/A_R) in conjunction with the effect of liquid velocity and a parameter,C ₀, describing the distribution of the liquid velocity and gas holdup across the riser on the bubble coalescence rates. The mixing performance of the reactor was dominated by the frequency of the passage of the broth through the end sections of the reactor. An optimum liquid height above the draft tube, for liquid mixing was demonstrated, above which no further improvement in mixing occurred. The liquid velocity and degree of gas entrainment showed little dependency on top section size for both sparger configurations. Extreme dissolved oxygen heterogeneity was demonstrated around the vessel with both sparger configurations and was shown to be detrimental to the oxygen uptake rate of the baker's yeast. Dissolved oxygen tensions below 1% air saturation occurred along the length of the riser and then rose in the downcomer. The greater oxygen transfer rate in the downcomer than in the riser was caused by the combined effects of a larger slip velocity in the downcomer which enhancedk _La and gas residence time, high downcomer gas holdup, and the change in bubble size distribution between the riser and downcomer. The position of greatest oxygen transfer rate in the downcomer was shown to be affected by the reactor from the influence on downcomer liquid linear velocity. UCL is the Biotechnology and Biological Sciences Research Council sponsored Advanced Centre for Biochemical Engineering and the Council's support is greatly acknowledged.     

Interpretation of gas residence time distributions in large airlift tower loop reactors

Gas-residence time distribution (RTD) response curves measured in a 23 m high pilot plant airlift tower loop reactor, which consisted of a riser, a special downcomer construction and at the top of the riser a large head. The measurements were evaluated by means of a deterministic dispersion model, which yielded the following particular parameters for the riser, downcomer and the head: Gas-Bo numbers, gas-mean residence times, gas holdups, liquid velocities, gas and liquid circulation times as well as a fraction of the large and small bubbles in a model medium (water) and during cultivation of baker's yeast.List of Symbols A cross section - Bo Bodenstein number - Bo _d (= l _d w _G,d /D _d ) - Bo _h (= l _h w _G,h /D _h ) - Bo _r (= l _r w _G,r /D _r ) - D longitudinal dispersion coefficient - E gas holdup - E(t) RTD-density function - L, l length parameter - q fraction of the gas throughput which is not recirculated (approximately equal to fraction of the large bubbles) - r fraction of the throughput which is recirculated (approximately equal to the fraction of the small bubbles) - t _c circulation time - t _cL liquid circulation time - t _c,L ^* liquid circulation time calculated from the measured w _Ld in the downcomer - V ^h hydrodynamical calculated gas-liquid volume - V _d ^h (=V _d+0.75/2 V _k ) - V _k ^h =(0.25V _k ) - V _r ^h = (V _r+0.75/2 V _k ) - V _L liquid volume - V _G dispersed gas volume - V _G ^* gas throughput at the gas distributor (given in m³/h) under standard conditions, 1 bar and 25°C) - V _G,d ^* gas throughput in downcomer (=V _G ^* ) - V _G,h ^* gas throughput in head (=V _G ^* ) - V _G,r ^* gas throughput in riser (V _G ^* (1+) - w _g gas velocity - w _G,rel relative gas velocity with respect to the liquid velocity w _L - w _G,d gas velocity in the downcomer (=w _G,rel –w _Ld ) - w _G,h gas velocity in the head (=w _G,rel ) (since wLh = o) - w _G,r gas velocity in the riser (=w _G,rel +w _Lr ) - w _L liquid velocity - w _L,d liquid velocity in the downcomer measured with mass flow meter - w _sg ·w _SL superficial gas and liquid velocities - first moment of the response curve - mean residence time Indices d downcomer - G gas phase - h head - L liquid phase - r riser - h hydrodynamic (upper position) Dedicated to the 65th birthday of Proffessor Fritz Wagner.The authors gratefully acknowledge the financial support by the Krupp Industrietechnik, Grevenbroich and the support of Pleser Co, Darmstadt. H. M. Rüffer thanks the Verband der Chemischen Industrie for a Fond der Chemie scholarship, and W. Liwei thanks the government of Lower Saxony for a graduate scholarship.     

Hydrodynamic characteristics and overall volumetric oxygen transfer coefficient of a new multi-environment bioreactor

The hydrodynamic characteristics and the overall volumetric oxygen transfer coefficient of a new multi-environment bioreactor which is an integrated part of a wastewater treatment system, called BioCAST, were studied. This bioreactor contains several zones with different environmental conditions including aerobic, microaerophilic and anoxic, designed to increase the contaminant removal capacity of the treatment system. The multi-environment bioreactor is designed based on the concept of airlift reactors where liquid is circulated through the zones with different environmental conditions. The presence of openings between the aerobic zone and the adjacent oxygen-depleted microaerophilic zone changes the hydrodynamic properties of this bioreactor compared to the conventional airlift designs. The impact of operating and process parameters, notably the hydraulic retention time (HRT) and superficial gas velocity (U _G), on the hydrodynamics and mass transfer characteristics of the system was examined. The results showed that liquid circulation velocity (V _L), gas holdup (ε) and overall volumetric oxygen transfer coefficient ( $ k_{\text{L}} a_{\text{L}} $ ) increase with the increase of superficial gas velocity (U _G), while the mean circulation time (t _c) decreases with the increase of superficial gas velocity. The mean circulation time between the aerobic zone (riser) and microaerophilic zone (downcomer) is a stronger function of the superficial gas velocity for the smaller openings (1/2 in.) between the two zones, while for the larger opening (1 in.) the mean circulation time is almost independent of U _G for U _G ≥ 0.023 m/s. The smaller openings between the two zones provide higher mass transfer coefficient and better zone generation which will contribute to improved performance of the system during treatment operations.     

Gas holdup and mass transfer characteristics of carboxymethyl cellulose solutions in a bubble column with a radial gas sparger

The gas phase holdup and mass transfer characteristics of carboxymethyl cellulose (CMC) solutions in a bubble column having a radial gas sparger have been determined and a new flow regime map has been proposed. The gas holdup increases with gas velocity in the bubbly flow regime, decreases in the churn-turbulent flow regime, and increases again in the slug flow regime. The volumetric mass transfer coefficient (k _La) significantly decreases with increasing liquid viscosity. The gas holdup and k _La values in the present bubble column of CMC solutions are found to be much higher than those in bubble columns or external-loop airlift columns with a plate-type sparger. The obtained gas phase holdup ( _g) and k _La data have been correlated with pertinent dimensionless groups in both the bubbly and the churn-turbulent flow regimes.List of Symbols a m^–1 specific gas-liquid interfacial area per total volume - A _d m² cross-sectional area of downcomer - A _r m² cross-sectional area of riser - d _b m individual bubble diameter - d _vs m Sauter mean bubble diameter - D _c m column diameter - D _L m²/s oxygen diffusivity in the liquid - Fr Froude number, U _g/(g D_c)^1/2 - g m/s² gravitational acceleration - G _a Galileo number, gD _c ^{3 2}/² _app - H _a m aerated liquid height - H _c m unaerated liquid height - K Pa · sⁿ fluid consistency index - k _L a s^–1 volumetric mass transfer coefficient - n flow behavior index - N _i number of bubbles having diameter d _bi - Sc Schmidt number, _app/( D _L) - Sh Sherwood number, k _L a D _c ² /D_L - U _sg m/s superficial gas velocity - U _gr m/s superficial riser gas velocity - V _a m³ aerated liquid volume - V _c m³ unaerated liquid volume - N/m surface tension of the liquid phase - _g gas holdup - _app Pa · s effective viscosity of non-Newtonian liquid - kg/m³ liquid density - ý s^–1 shear rate - Pa shear stress     

Flow characteristics of a pilot-scale airlift fermentor

The effects of aeration on the flow characteristics of water in a glass pilot-scale airlift fermentor have been examined. The 55-L capacity fermentor consisted of a 15.2-cm-i.d. riser column with a 5.1-cm-i.d. downcomer tube. It was found that the average bubble size diminished with increased aeration. Typically, average bubble sizes ranged from 4.32 mm at a superficial gas velocity of 0.64 cm/s to 1.92 mm at 10.3 cm/s. A gas holdup of 0.19 was attained with superficial gas velocities (v_s) on the order of 10 cm/s, indicating the highly gassed nature of the fluid in the riser section of the fermentor. Circulation velocities of markers placed in the fermentor decreased with increasing aeration rates due to increased turbulence and axial liquid back mixing within the riser section. Actual volumetric liquid circulation rates remained relatively constant (0.36–0.49 L/s) for values of (v_s) up to 10 cm/s. Based on theoretical calculations, the ascending velocity of bubbles in a swarm reached 54 cm/s in the range of (v_s) values studied.     

Oxygen transfer measurements during yeast fermentations in a pilot scale airlift fermenter

Measurements of oxygen transfer were made during cultivation of the yeast Saccharomyces cerevisiae in a 90–250 litre working volume concentric tube airlift fermenter. Results demonstrated that the rate of oxygen transfer varies with position in the fermenter, being higher in the riser and top-section than in the downcomer and lowest near the base of the fermenter. The time for liquid circulation was generally smaller than the time constant for oxygen transfer (1/k_La) indicating that the rate of oxygen transfer was slow compared to the rate of liquid movement. Measured dissolved oxygen concentrations therefore did not represent the equilibrium arising from the balance between the rates of oxygen transfer and oxygen depletion. Hence measuredk _L a values were not representative of local oxygen transfer conditions but instead were indicators of the rate of mass transfer the liquid flow had encountered prior to reaching the point of measurement. Generally the individual rates of oxygen transfer in the vessel were found to increase with increasing vessel height.     

Mathematical models for mixing in deep jet bioreactors: analysis

A mathematical model for single and multi step deep-jet bioreactors is presented. A stagewise approach based on macroscopic mechanistic model which divides the reactor into compartments with good quality of mixing and plug flow regions (macromixer), was used. For the mathematical representation of this model a system of differential equations, describing the concentration of tracer in structural elements based on mass balance, and the Runge-Kutta-Fehlberg numerical method of integration, was applied. The mixing time in a 300 dm³ tank was determined by conductivity method with NaCl as tracer.List of Symbols V _g dm³ total volume of liquid - V ₁; V ₆ dm³ volumes of ideally mixed compartments in the vessel - V ₂; V ₇ dm³ volumes of macromixer in the inner circulation flows - V ₃; V ₉ dm³ volumes of liquid phase in the pump - V ₄; V ₈ dm³ volumes of liquid phase in the pipe between the vessel and the pump - V ₅; V ₁₀ dm³ volumes of liquid phase in pipes between the pump and the air input system, including falling jet - F _E; F _E,1; F _E,2 dm³/s the inner volumetric circulation flow rates accross the macromixers - F _E,3; F _E,4 dm³/s exchanges volumetric flow rates between two ideally mixed compartments in the vessel - F _cir; F _1,cir; F _2,cir dm³/s external volumetric circulation flow rates (pumping capacity) - t _A s time interval of puls application - t _AA s time point of impuls application related to the free chosen point of simulation - t _end s end time of simulation - F _qu g²/dm⁶ sum of quadratic error - C _*,* kg/m³ concentration of the 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 full mixed tracer - i index of an arbitrary tank - C _sim kg/m³ calculated concentration of the tracer by numerical integration method     

Gas hold-up in converging-diverging tube airlift fermenter

Summary Fractional gas holdup study was carried out in two airlift fermenters: one having of conventional design, the other having an asymetric riser arm. Air flow rate was varied from 1.5 to 9.0 cm/sec and gas hold-up values compared. Fractional gas holdup, _G, was strongly dependent on superficial gas velocity and initial liquid height. The modified fermenter always showed a higher gas holdup than the conventionally designed one.Symbols ALF Airlift Fermenter - CDT Convergent-divergent Tube - UT Uniform Tube - U_G Superficial gas velocity, cm/s - h_i Initial liquid height in riser, cm - H_i Dispersed liquid height in riser, cm - H_O Dispersed liquid height in downcomer, cm - K,m,n Constant - a,a Constant - A_d Riser cross sectional area, cm^z - A_r Downcomer cross sectional area, cm^z - U_b Bubble rise velocity, cm/s - g Acceleration due to gravity, cm/s^z - d_B Bubble diameter, cm - R_ev Bubble's Reynolds number, dimensionless Greek Letters _G Fractional gas holdup, dimensionless - {ITG9}{INL} Liquid density, g/cc - {IT}{INL} Liquid viscosity, poise(g/cm.s) - {ITGS}{INL} Liquid surface tension, dyne/cm - porous plate pore diameter, cm     

A novel split-cylinder airlift reactor for fed-batch cultures

Conventional airlift reactors are not adequate to carry out variable volume processes since it is not possible to achieve a proper liquid circulation in these reactors until the liquid height is higher than that of the downcomer. To carry out processes of variable volume, the use of a split-cylinder airlift reactor is proposed, in the interior of which two multi-perforated vertical baffles are installed in order to provide several points of communication between the reactor riser and downcomer. This improves the liquid circulation and mixing at any liquid volume. In fed-batch cultures, it is important to know how liquid height affects the hydrodynamic characteristics and the volumetric oxygen transfer coefficient since this impacts on the kinetic behavior of any fermentation. Thus, in the present work, the effect of the liquid height on the mixing time, the overall gas hold-up, and the volumetric oxygen transfer coefficient of the proposed airlift reactor were determined. The mixing time was increased and the volumetric oxygen transfer coefficient decreased due to the increase of the liquid height in the reactor in all the superficial gas velocities tested, which corresponded to a pseudohomogeneous flow regime. The experimental values of the mixing time and the mass-transfer coefficient were properly described through correlations in which the U_GR/H_L ratio was used as the independent variable. Thus, this variable might be used to describe the hydrodynamic behavior and the oxygen transfer coefficient of airlift reactors when such reactors are used in processes where the liquid volume changes with time. However, these correlations are useful for the particular device and for the particular operating conditions at which they were obtained. These empirical correlations are useful to understand some factors that influence the mixing time and volumetric oxygen transfer coefficient, but such correlations do not have a sufficient predictive potential for a satisfactory reactor design. The overall gas hold-up values were not significantly affected when the liquid height was lower than the downcomer height. However, the values decreased abruptly when the reactor was operated with liquid heights over the downcomer height, especially at high superficial gas velocities.     

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