Evolution axiale du coefficient de transfert de matière en colonne à bulles et en lit fluidisé triphasique

Liquid phase volumetric mass transfer coefficients for oxygen are determined in a three-phase fluidized bed and in a bubble column. The concept of exponential decreasing axial variation of volumetric mass transfer coefficient leads to a better representation of oxygen concentration profiles inside the column. Compared to the bubble column, k_la axial variations are more important in the lower part of the fluidized bed column, where solid particles increase the coalescence phenomenum, particularly with viscous liquids.     

Gas backmixing in the dense region of a circulating fluidized bed

The gas backmixing characteristics in a circulating fluidized bed (0.1 m-IDx5.3-m high) have been determined. The gas backmixing coefficient (D_ba) from the axial dispersion model in a low velocity fluidization region increases with increasing gas velocity. The effect of gas velocity onD _ba in the bubbling bed is more pronounced compared to that in the Circulating Fluidized Bed (CFB). In the dense region of a CFB, the two-phase model is proposed to calculate D_bc from the two-phase model and mass transfer coefficient (k) between the crowd phase and dispersed phase. The gas backmixing coefficient and the mass transfer coefficient between the two phases increase with increasing the ratio of average particle to gas velocities (U_p/U_g).     

Bubble breakage phenomena,phase holdups and mass transfer in three-phase fluidized beds with floating bubble breakers

The individual phase holdups and mass transfer characteristics in three-phase fluidized beds with different floating bubble breakers have been determined in a 2.0 m high Plexiglas column of inner diameter 0.142 m. The bubble breaking phenomena by the breakers have been studied via a photographic method in a two-dimensional Plexiglas column. The volumetric mass transfer coefficient k_La in three-phase fluidized beds with hexagonal-shaped breakers is up to 40% greater than that in beds without floating bubble breakers. The bed porosity ε_L + ε_g, gas-phase holdup ε_g, and volumetric mass transfer coefficient k_La increase with an increase in the volume ratio of floating bubble breakers to solid particles, V_f/V_s, up to around 0.15, and thereafter decrease with V_f/V_s in three-phase fluidized beds with floating bubble breakers. Also, k_La increases with increasing breaker density, projected area and contact angle between the floating bubble breakers and the water. The volumetric mass transfer coefficients in terms of the Sherwood number in three-phase fluidized beds with the various floating bubble breakers have been correlated with the volume ratio of floating bubble breakers to solid particles, the particle Reynolds number based on the local isotropic turbulence theory and the modified Weber number.     

Modelling turbulent fluidized bed reactors: tracer and fibre optic probe studies

Phenomenological models for turbulent fluidized beds are presented in this study. These models are based on a “core-annulus” representation of the turbulent fluidized bed.Three flow regions are considered: (1) gas flows through a dense annular region and is either perfectly mixed or in plug flow; (2) gas circulates in the core as bubbles in plug flow; (3) gas is perfectly mixed in a dense emulsion phase, also in the core zone. The models also account for mass transfer between different regions by assuming various possible gas exchange paths.A new technique which combines novel reflective fibre optic probes and statistical signal treatment is used to measure local flow properties. Results from fibre optic experiments coupled with those from an inert non-adsorbing tracer (helium) allow for mass transfer parameter assessment.These data demonstrate the importance of incorporating an annular region in the simulation of the main bed section of turbulent fluidized beds. Modelling results of this work strongly suggest the critical importance of gas exchange between bubbles in the core and a pseudo-homogeneous annular region.     

Effect of bubble interaction on interphase mass transfer in gas fluidized beds

A non-interfering technique has been used to measure the concentration of ozone in pairs of bubbles injected into a bed of inactive 390 μm glass beads fluidized by ozone-free air. The transfer of the ozone tracer from the bubble phase to the dense phase is enhanced when compared to the transfer from isolated bubbles in the same particles and the same column. Bubble growth is also greater for the case where pairs of bubbles are introduced than when bubbles are present in isolation. Enhancement of interphase mass transfer for interacting bubbles in the present work and in previous studies incr with particle size and can be explained in terms of enhancement of the throughflow (or convective) component of transfer while the diffusive component unaltered. This mechanism leads to new equations for estimating interphase mass transfer in freely bubbling fluidized beds.     

Gas-Liquid mass transfer in a three-phase fluidized bed with floating bubble breakers

The effects of liquid and gas velocities, particle size and volume ratio of floating bubble breakers to solid particles (V_f/V_s) on both the volumetric mass transfer coefficient, k_la, and the gas-liquid interfacial area, a, have been determined in three-phase fluidized beds with floating bubble breakers. Beds having a volume ratio (V_f/V_s) of about 0.15 showed a maximum increase in both k_la and a of about 30% in comparison to that in the corresponding bed without floating bubble breakers. The volumetric mass transfer coefficient in three-phase fluidized beds with or without floating bubble breakers can be estimated from the surface renewal frequency of liquid microeddies and the particle size.     

Gas transfer from bubbles in a fluidized bed to the dense phase — II. Experiments

An experimental check was made upon the theory given in Part I. Cracking catalyst was used as a solid and differently adsorbed tracer gases were used. In a two-dimensional fluidized bed bubbles were formed underneath a gauze cap, while solid flowed along the bubble at the corresponding bubble velocity. Tracer injections provided the value for the transfer coefficient. In three-dimensional beds of 18 and 90 cm dia. large traced gas bubbles were injected. Tracer concentration was detected at certain heights. From the decrease the transfer coefficient was calculated. In the 90 cm bed the transfer coefficient was also calculated from residence time distribution measurements when the dense phase was perfectly mixed.It shows, that the two-dimensional bubble confirms the theory. For three-dimensional bubbles the transfer is higher than theoretically predicted, especially when the dense phase is expanded.     

Numerical simulations of gas-liquid mass transfer in bubble columns with a CFD-PBM coupled model

Gas-liquid mass transfer in a bubble column in both the homogeneous and heterogeneous flow regimes was studied by numerical simulations with a CFD-PBM (computation fluid dynamics-population balance model) coupled model and a gas-liquid mass transfer model. In the CFD-PBM coupled model, the gas-liquid interfacial area a is calculated from the gas holdup and bubble size distribution. In this work, multiple mechanisms for bubble coalescence, including coalescence due to turbulent eddies, different bubble rise velocities and bubble wake entrainment, and for bubble breakup due to eddy collision and instability of large bubbles were considered. Previous studies show that these considerations are crucial for proper predictions of both the homogenous and the heterogeneous flow regimes. Many parameters may affect the mass transfer coefficient, including the bubble size distribution, bubble slip velocity, turbulent energy dissipation rate and bubble coalescence and breakup. These complex factors were quantitatively counted in the CFD-PBM coupled model. For the mass transfer coefficient k_l, two typical models were compared, namely the eddy cell model in which k_l depends on the turbulent energy dissipation rate, and the slip penetration model in which k_l depends on the bubble size and bubble slip velocity. Reasonable predictions of k_la were obtained with both models in a wide range of superficial gas velocity, with only a slight modification of the model constants. The simulation results show that CFD-PBM coupled model is an efficient method for predicting the hydrodynamics, bubble size distribution, interfacial area and gas-liquid mass transfer rate in a bubble column.     

Bubble shape,gas flow and gas–liquid mass transfer in pulp fibre suspensions

Gas–liquid mass transfer in pulp fibre suspensions in a batch‐operated bubble column is explained by observations of bubble size and shape made in a 2D column. Two pulp fibre suspensions (hardwood and softwood kraft) were studied over a range of suspension mass concentrations and gas flow rates. For a given gas flow rate, bubble size was found to increase as suspension concentration increased, moving from smaller spherical/elliptical bubbles to larger spherical‐capped/dimpled‐elliptical bubbles. At relatively low mass concentrations (C_m = 2–3% for the softwood and C_m ? 7% for the hardwood pulp) distinct bubbles were no longer observed in the suspension. Instead, a network of channels formed through which gas flowed. In the bubble column, the volumetric gas–liquid mass transfer rate, k_La, decreased with increasing suspension concentration. From the 2D studies, this occurred as bubble size and rise velocity increased, which would decrease overall bubble surface area and gas holdup in the column. A minimum in k_La occurred between C_m = 2% and 4% which depended on pulp type and was reached near the mass concentration where the flow channels first formed.     

Improving Gas‐Liquid Mass Transfer in Bubble Columns by Applying Low‐Frequency Vibrations

We show that application of low‐frequency vibrations, in the 50–200 Hz range, to the liquid phase of an air‐water bubble column causes significantly smaller bubbles to be generated at the distributor plate. For bubble column operation in the homogeneous flow regime, measurements of the volumetric mass transfer coefficient using the oxygen absorption technique show that the increase in the k_La values ranges from 50–100 % depending on the flow rate. It is concluded that application of low‐frequency vibration has the potential of improving the performance of bubble columns.     

Analogous description of the hydrodynamics of gas-solid fluidized beds and bubble columns

In this paper we stress analogies in the hydrodynamic behaviour of gassolid fluidized beds and bubble columns. Using published experimental data, it is demonstrated that the analogous hydrodynamic-behaviour is not only qualitative but also quantitative in nature. Specifically, we show the following.(1) The gas holdup in the homogeneous regimes of bubble columns and fluidized beds can be modelled in a unified way using V_slip = υ_∞(1 − ϵ_d)ⁿ⁻¹, where V_slip refers to the slip velocity between the dispersed (bubbles or particles) and continuous phases and ϵ_d the dispersed phase holdup. The Richardson-Zaki exponent n decreases with increasing gas density.(2) The transition from homogeneous to heterogeneous flow regimes in gasliquid bubble columns and gassolid fluid beds is delayed by increasing system pressure. Extrapolation of the influence of increased gas density allows us to consider liquidliquid dispersions and liquidsolid fluid beds as limiting cases.(3) In the heterogeneous flow regime of operation the classic two-phase theory of fluidized beds can be applied with profit to also describe the hydrodynamics of gasliquid bubble columns provided that the “dilute” phase is identified with the fast-rising large bubbles and the “dense” phase is identified with the liquid phase containing entrained “small” bubbles. Tentative analogies can also be drawn for the interphase mass transfer processes.(4) The “dense” phase backmixing can be modelled in a unified manner.(5) The two-phase theory can be extended to describe slurry reactors.It is argued that, because of cross-fertilization of concepts and information, appreciation of analogies can be invaluable tool in scaling up.     

Influence of solids concentration and static mixers on fluid dynamics in three-phase fluidized bed bioreactors

The influence of solids concentration and static mixers on the hydrodynamics of the gas phase was studied for a three-phase fluidized bed bioreactor (air, nutrient solution, biocatalyst Ca alginate beads). Axial gas hold-up profiles, radial gas velocity profiles, mean bubble diameter and gas/liquid interfacial area per unit volume were measured in a bubble column (D_R = 0.142 m, H_R = 1.748 m). The influence of solids concentration on the gas hold-up is insignificant; static mixers enhance the gas hold-up in the reactor volume element in which they are installed. Axial gas velocity decreases with increasing solids concentration. At high solids concentrations, static mixers exert little influence on the gas phase but, at low concentrations, they do. A model is suggested to describe the influence of solids concentration (characterized by turbulent viscosity v_t) and static mixers (characterized by profile parameter n) on the gas velocity profile.     

Population balance modelling of bubbling fluidized bed reactors—I. well-stirred dense phase

A detailed population balance model is presented for a fluidized bed reactor incorporating: the formation of bubbles at the grid plate, their rise with velocities governed by their sizes, random coalescence between bubbles, gas exchange between bubbles and the dense phase, and a first order chemical reaction in the dense phase under well-mixed conditions. Reaction conversion is calculated as a function of dimensionless parameters relating the rates of various competing processes such as coalescence, dense phase mixing, mass exchange between bubbles and dense phase and reaction rate. Comparison of conversions with those of Davidson et al. (1977) show significant variations indicating that the dynamics of bubble size distributions could have non-trivial effects on the extent of reaction. Fluctuations in bubble populations did not seem strong enough to translate to strong fluctuations in reaction conversion.     

Heat transfer between an isolated bubble and the dense phase in a fluidized bed

Heat transfer between the bubble and dense phases of a bubbling fluidized bed plays a very important role in the system performance, especially for applications involving solids drying and gas‐phase combustion. However, very few experimental data are available on this subject in the literature. An experimental and modelling investigation on the heat transfer behaviour of isolated bubbles injected into an incipiently fluidized bed is reported in this paper. A new single‐thermocouple technique was developed to measure the heat transfer coefficient. The effects of bed particle type and size, and bubble size on the heat transfer coefficient were examined. The heat transfer coefficient was found to exhibit a maximum as the bubble size increased in the bubble size range investigated. The bed particle size had a comparatively small effect on the heat transfer coefficient. A simple mathematical model was developed which provides good agreement with experimental data.     

Tomographic imaging of a conical fluidized bed of dry pharmaceutical granule

Hydrodynamics in a conical fluidized bed were studied using electrical capacitance tomography (ECT) for a bimodal and mono-disperse particle size distribution (PSD) of dry pharmaceutical granule. The bimodal PSD exhibited a continuous distribution with modes at 168 and 1288 μm and contained approximately 46% Geldart A, 32% Geldart B and 22% Geldart D particles by mass. The mono-disperse PSD had a mean particle size of 237 μm and contained approximately 71% Geldart A, 27% Geldart B, and 2% Geldart C particles by mass. The granule particle density was 830 kg/m³. Experiments were conducted at a static bed height of 0.16 m for gas superficial velocities ranging from 0.25 to 2.50 m/s for the mono-disperse PSD, and from 0.50 to 3.00 m/s for the bimodal PSD. These gas velocities covered both the bubbling and turbulent fluidization regimes. An ‘M’-shaped time-averaged radial voidage profile appeared upon transition from bubbling to turbulent fluidization. The ‘M’-shaped voidage profile was characterized by a dense region near the wall of the fluidized bed with decreasing solids concentration towards the centre. An increased solids concentration was observed in the middle of the bed. Frame-by-frame analysis of the images showed two predominant bubble types: spherical bubbles with particle penetration in the nose which created a core of particles that extended into, but not through, the bubble; and spherical bubbles. Penetrated bubbles, responsible for the ‘M’ profile, were a precursor to bubble splitting; which became increasingly prevalent in the turbulent regime.     

EFFECT OF STATIC LIQUID HEIGHT ON GAS-LIQUID MASS TRANSFER IN A DRAFT-TUBE BUBBLE COLUMN AND THREE-PHASE FLUIDIZED BED

Experiments are performed under batch-liquid operating conditions to investigate the effect of static liquid height on the gas-liquid mass transfer coefficient (K_La) in a draft-tube bubble column (DTBC) and a draft-tube three-phase fluidized bed (DTFB). In addition, the effects of column diameter, gas-distributor, and draft-tube diameter are studied. The results indicate that for a given system with a porous plate gas-distributor at low superficial gas velocities (<70 m/hr), increasing static liquid height decreases K_La. At high gas velocities, K_La is independent of the static liquid height. For systems with a perforated gas-distributor, there is no effect of static liquid height on K_La. The formation of small dispersed bubbles at low gas velocities in the porous plate distributor system accounts for the considerably high K_La values and the observed effect of liquid height. On the other hand, the formation of large spherical-cap bubbles and the bubble coalescence at high gas velocities reduce the performance of the porous plate distributor system to that of the perforated one.     

Kinetic simulation of a compact reactor system for hydrogen production by steam reforming of higher hydrocarbons

A bubbling fluidized bed membrane reactor for steam reforming of higher hydrocarbons is modelled, using n‐heptane as a model component to represent steam reforming of naphtha. The reformer is modelled as a bubbling fluidized bed reactor, consisting of two pseudo phases, a dense phase and a bubble phase, both in plug flow. In situ H₂ permselective membranes remove H₂ continuously as a pure product, greatly enhancing the H₂ yield per mole of heptane fed. A fluidized bed membrane reformer for higher hydrocarbons could give a very compact reactor system combining all the units from the pre‐reformer to the hydrogen purification system in a traditional steam reforming plant into a single unit.     

Type transition in onset condition of turbulent fluidization

The type transition in onset condition of turbulent fluidization in gas fluidized beds was investigated to obtain the relation representing more precise roles of physical properties of gas and solid particles. The type transition in onset condition of turbulent fluidization occurs at Archimedes number of 20.87 by type transition of bubble breakup. The maximum stable bubble diameter (d _bmax ) is greater than the equilibrium bubble diameter (d _beq ) in the region, Ar< 20.87, but d_beq>d _bmax in the region, Ar>20.87. Therefore, the onset of turbulent fluidization is determined in the region, Ar<20.87, by d _beq and in the region, Ar>20.87, by d _bmax as the limit of bubble growth. The uc decreases in the region, Ar<20.87, but increases in the region, Ar>20.87 as temperature increases.     

Expansion characteristics of an anaerobic fluidized bed reactor with internal biogas production

More realistic dynamic bed‐expansion experiments using a three‐phase anaerobic fluidized bed reactor (AFBR) with and without internal biogas production were conducted for the establishment of correlation equations for the mean volume ratio of wakes to bubbles (k). A predictive model was also developed for the expansion characteristics of the three‐phase AFBR with internal biogas production. The predicted bed‐expansion heights (H_GLS) deviated by only ±10% from the experimental measurements for the three‐phase AFBR. According to the modeling results, if a three‐phase AFBR is loaded into a carrier with low specific gravity (dry density of carrier, ρ_md = 1.37 g cm^?3; wet density of carrier, ρ_mw = 1.57 g cm^?3) and operated at a high superficial liquid velocity (u_l = 4.0 cm s^?1), the ratio of H_GLS to H_LS at a high superficial gas velocity (u_g = 1.5 cm s^?1) can reach as high as 271%. A higher fluidized‐bed height has a greater effect on the bed‐expansion behavior because of the decrease in liquid pressure (surrounding gas bubbles) along the fluidized‐bed height. From parametric sensitivity analyses, H_GLS is most sensitive to the parameter reactor width (X), especially within a small ΔX/X₀ range of ±10%; sensitive to ρ_mw, diameter of the carrier, ρ_md and total mass of carrier and least sensitive to u_l, biofilm thickness and u_g. Copyright © 2005 Society of Chemical Industry