MASS TRANSFER AND PHASE HOLDUP CHARACTERISTICS IN THREE-PHASE FLUIDIZED BEDS

The effects of liquid surface tension (42.6 ∼ 72,4 mN/m) and viscosity (1 ∼214mPa • sⁿ), liquid (0.01 ∼0.12m/s) and gas (0.01 ∼0.20m/s) velocities and particle sizes (1 — 8 mm) on phase holdup and mass transfer coefficient ( k_La) have been determined in a 0.142 m-I.D. × 2.0 m-high Plexiglas column. The gas phase holdup increases with liquid velocity, and the rate of increase in gas phase holdup sharply increases with gas velocity in the bed of surfactant solutions. In the beds of 1.0 and 1.7 mm glass beads, the bed contraction occurs whereas in the beds of 2.3 mm glass beads the bed contraction does not occur with an aqueous soltuion of ethanol (σ = 50.4 mN/m). The value of k_La increases with decreasing surface tension (σ ) but it decreases exponentially with increasing liquid viscosity in continuous bubble columns and three-phase fluidized beds. In three-phase fluidized beds with surfactant solutions, k_La increases with gas and liquid velocities and particle size. In three-phase fluidized beds of viscous or surfactant soltuions, k_L,a can be estimated in terms of the energy dissipation rate based on the isotropic turbulence theory and a flow regime map is proposed based on the drift flux theory.     

气-液-固三相逆流化床热量传递研究

在内径0.152 m,高2.5 m的气-液-固三相逆流化床中系统研究了热量传递特性特性。获得了气体和液体速度及聚乙烯和聚丙烯颗粒密度对内置加热器与床层间热量传递系数的影响规律。研究结果表明密度相对高的聚乙烯颗粒的逆流化床的热量传递系数比密度相对低的聚丙烯颗粒的逆流化床的热量传递系数大;随着气体速度的增加,热量传递系数增加。然而,随着液体速度增加,热量传递系数具有最大值。在热量传递系数达到最大值时对应的液体速度随着颗粒密度或气体速度的增加而降低。     

Bubble drift velocity from the bed collapse technique in three-phase fluidized beds

Transient behavior of a bed collapsing after cut-off of gas supply into a three-phase fluidized bed was determined in a 0.21 m-diameter half-tube acrylic column having a test section 1.8 m high. The transient behavior of the bed collapse after cut-off of the gas supply to the beds was monitored by a video camera (30 frames/s). A theory was developed to account for the dynamic behavior of the bed collapse after the gas supply shut-off to three-phase fluidized beds. The bubble drift velocity was theoretically calculated by gas and liquid phase holdups at steady state condition. At a liquid velocity of 0.103 m/s and gas velocity of 0–0.023 m/s, bubble size was uniform in the dispersed bubble flow regime. However, as the gas velocity increased above 0.023 m/s, the discrete or coalesced bubble flow regime could be observed. The agreement between the predicted and experimental values is acceptable in the dispersed bubble flow regime, but the agreement becomes poorer with increasing gas velocity.     

A highly elevated mass transfer rate process for three-phase,liquid-continuous fluidized beds

Average gas holdup and gas-to-liquid mass transfer in three-phase fluidized beds with non-Newtonian fluids were studied. The effects of liquid property, gas distributor type and magnetic field intensity on mass transfer coefficient and overall gas holdup were examined. The volumetric gas-to-liquid mass transfer coefficient was determined by fitting the oxygen concentration profile data across the bed to the axial dispersion model. The average gas holdup and mass transfer coefficient were all correlated with operating parameters including gas velocity and effective viscosity.Experimental results showed that a three-fold increase in mass transfer coefficient and a two-fold increase in average gas holdup were observed with properly designed liquid property and gas distributor. A modified process was developed to highly elevate the volumetric gas-to-liquid mass transfer rate. The bubble coalescing property of three-phase fluidized beds with small particles is eliminated, and its application to biotechnology and enzyme-catalyzed processes with high gas-to-liquid mass transfer rate could be achieved.     

Performance of a three-phase fluidized bed as a reactive distillation device

A new reactive distillation device, the multistage gas/liquid/solid three-phase fluidized bed, has been developed. The flow regimes of the multistage three-phase fluidized bed have been studied and the regimes can be divided into the liquid leakage regime, the dispersed bubble regime, and the coalesced bubble regime. Liquid velocity has a much smaller effect on phase holdups in this device than in conventional three-phase fluidized beds. The three phase fluidized bed is used as a reactive distillation device for the hydrolysis of methyl acetate. Much higher reaction conversion than the equilibrium value and high catalyst-contacting efficiency are obtained. Different methods of feeding the water into the reactive distillation section are studied.     

Pressure fluctuations and bubble size in viscous three-phase circulation fluidized bed bioreactors

Characteristics of pressure fluctuations and bubble size were investigated in the riser of a three-phase circulation fluidized bed bioreactor with viscous liquid medium, whose diameter is 0.102 m (ID) and 3.5 m in height. Effects of gas (0.01–0.07 m/s) and liquid (0.17–0.23 m/s) velocities and liquid viscosity (0.96–38 mPa·s) on the bubble size in the riser were examined. The bubbling phenomena in the bioreactor with viscous liquid medium were interpreted effectively by measuring and analyzing the pressure fluctuations by adopting chaos theory. The bubble size increased with increasing gas velocity or liquid viscosity, but decreased with increasing liquid velocity. The bubbling phenomena became more complicated and bubble size distribution tended to broad, with increasing gas velocity or liquid viscosity. The bubble size was well correlated in terms of correlation dimension of pressure fluctuations as well as dimensionless groups within these experimental conditions.     

Effect of pseudoplastic behaviour of liquid in co-current three-phase fluidized beds on bed expansion

Gas hold-up and bed expansion measurements were carried out for a bed of glass beads fluidized in Newtonian liquids and non-Newtonian liquids with gas. The value of gas hold-up increased and decreased with increasing particle size and liquid velocity, respectively. The effect of rheological properties on gas hold-up was insignificant and therefore the gas hold-up data for both Newtonian and non-Newtonian fluids were reasonably fitted by the available correlation which had no liquid viscosity term. The bed voidage increased with increasing superficial liquid velocities and superficial gas velocities. The increase of the viscous non-Newtonian flow behaviours resulted in an increase of the bed voidage. The correlation for the bed voidage in three-phase fluidized beds was developed for gas-Newtonian or non-Newtonian liquid–solid three-phase systems by combining the generalized wake model and the correlation for liquid–solid two-phase systems proposed previously by the authors. The predictions for bed voidage were in reasonable agreement with the present experimental data for three-phase systems with Newtonian and non-Newtonian liquids in a wide range of Reynolds numbers.     

Axial dispersion characteristics in three phase fluidized beds

Axial dispersion coefficients in three-phase fluidized beds have been measured in a 0.152 m-ID x 1.8 m high column by the two points measuring technique with the axially dispersed plug flow model. The effects of liquid velocity (0.05–0.13 m/s), gas velocity (0.02–0.16 m/s) and particle size (3-8 mm) on the axial dispersion coefficient at the different axial positions (0.06–0.46 m) in the bed have been determined. The axial dispersion coefficient increases with increasing gas velocity but it decreases with an increase in particle size and exhibits a maximum value with an increase in the axial position from the distributor. The axial dispersion coefficients in terms of the Peclet number have been correlated in terms of the ratio of fluid velocities, the ratio of the panicle size to column diameter, and the dimensionless axial position in the bed based on the isotropic turbulence theory.     

Mass transfer in two-and three-phase fluidized beds

The effects of liquid (0.03-0.12 m/s) and as (0.04-0.20 m/s) velocities, and particle size (0-8.0 mm) on the volumetric mass transfer coefficients at the grid zone have been determined in a 0.152 mI.D. x 1.8 m high Plexiglas column. The volumetric mass transfer coefficient in the grid zone increases with increasing gas velocity and particle size. However, the coefficient exhibits a maximum value at an optimum bed porosity condition. The volumetric mass transfer coefficients in terms of the Sherwood number in three-phase fluidized beds have been correlated with the Schmidt number and particle Reynolds number which is related to the energy dissipation rate in the beds based on the local isotropic turbulence theory. Also, the coefficient has been correlated with the experimental variables.     

Particle dispersion in viscous three-phase inverse fluidized beds

Dispersion characteristics of low density fluidized particles such as polyethylene and polypropylene were investigated by using the stochastic method in three-phase inverse fluidized beds with viscous liquid medium ( in height). To establish the relationship between the pressure drop variation and the particle dispersion in test section, the histogram of pressure drop fluctuations were also measured and analyzed. Effects of operating variables such as gas and liquid velocities, liquid viscosity and media particle kind (density) on the fluctuating frequency, dispersion coefficient and exiting rate of media particles from the test section were determined. The fluctuating frequency and dispersion coefficient of particles increased with increasing gas or liquid velocity, but decreased considerably with increasing liquid viscosity in three-phase inverse fluidized beds. The dispersion coefficient of media particles of relatively higher density exhibited a value higher than that of lower density particles. The dispersion coefficients of particles were well correlated with operating variables in terms of dimensionless groups.     

THREE-PHASE FLUIDIZED BEDS WITH VISCOUS LIQUID: HYDRODYNAMICS AND MASS TRANSFER†

Hydrodynamics and gas/fiquid mass transfer in fluidized beds of glass spheres (3-8 mm diameter)were studied employing viscous aqueous solutions (16-53 mPas) Increasing liquid viscosity reduced the bubble disintegration capability of the particle beds. The most pronounced consequence was a strong decrease in the volumetric mass transfer coefficients (k_La) From a comparison of k_La in Newtonian and pseudoplastic liquids it is concluded that the effective shear rates in three-phase fluidized beds are higher than in bubble columns.     

Effect of air distributor on the fluidization characteristics in conical gas fluidized beds

The effect of an air distributor on the fluidization characteristics of 1 mm glass beads has been determined in a conical gas fluidized bed (0.1 m-inlet diameter and 0.6 m in height) with an apex angle of 20‡. To determine the effect of distributor geometry, five different perforated distributors were employed (the opening fraction of 0.009–0.037, different hole size, and number). The differential bed pressure drop increases with increasing gas velocity, and it goes from zero to a maximum value with increasing or decreasing gas velocity. From the differential bed pressure drop profiles with the distributors having different opening fractions, demarcation velocities of the minimum and maximum velocities of the partial fluidization, full fluidization, partial defluidization and the full defluidization are determined. Also, bubble frequencies in the conical gas fluidized beds were measured by an optical probe. In the conical bed, the gas velocity at which the maximum bed pressure drop attained increases with increasing the opening fraction of distributors.     

三相循环流化床中气泡大小及其分布的实验研究

用光纤探头技术对三相循环流化床中的气泡大小及其分布进行了系统研究 ,实验测定了操作条件对气泡大小及其分布的影响规律 .实验结果表明 ,三相循环流化床中气泡的大小分布可用对数正态分布表征 ,在实验条件下气泡平均直径在床中心区域较小且沿半径方向由中心向边壁逐渐增大 ,并随表观气速的增大而减小 ,随固含率的增大而增大 ,表观液速对气泡平均直径的影响较小     

Flow regimes in a gas–liquid–solid three-phase moving bed

The gas–liquid–solid three-phase moving beds could supply a potential solution for multiphase reactions with catalyst easily deactivated, and the flow regimes in it were studied by optical method and pressure drop measurement. Results showed that taking the trickle flow as the initial flow regime, the flow channels were more obvious as the particle velocity increased. When the initial flow regimes were pulse flow and bubble flow respectively, the pulse-to-trickle and bubble-to-pulse flow transitions mainly occurred at moderate-to-high particle velocities (0.01–0.04 m s⁻¹ under conditions used in this work). Moreover, the flow regime map in the three-phase moving bed was constructed and shown that the region of trickle flow increased and the region of bubble flow decreased. Finally, the application of three-phase moving beds was discussed, and it could be suitable for those reactions, which had to operate in the pulse flow, bubble flow, and transition zone.     

Transition from Low Velocity to High Velocity in a Three Phase Fluidized Bed

The onset liquid velocity demarcating the conventional and the circulating fluidization regimes of three‐phase fluidized beds was determined by measuring the time required to empty all particles in a batch fluidized bed at various liquid and gas velocities. Experiments were performed in a gas‐liquid‐solid circulating fluidized bed of 2.7 m in height using glass beads of 0.508 mm in diameter as solid phase and air and tap water as the fluidizing gas and liquid, respectively. The results show that gas velocity is a strong factor on the onset liquid velocity. Higher gas velocity yields a lower onset liquid velocity. It is also demonstrated that the onset liquid velocity has the same value as particle terminal velocity in a gas‐liquid mixture. Within the gas‐liquid‐solid circulating fluidization regime, the solids circulation rate is increased with the total liquid velocity and the auxiliary liquid velocity.     

Bubble properties and pressure fluctuations in bubble columns

The effects of gas (0.02-0.1 m/s) and liquid velocities (0.0-0.10 m/s) on the bubble properties and pressure fluctuations have been determined in a 0.376 m-IDx 2.1 m-high bubble column. The pressure fluctuations have been analyzed by resorting to the Fractal analysis; the time series of pressure fluctuation signals have been analyzed by means of the Rescaled range analysis and the Hurst exponent has been obtained. The bubble chord length and its rising velocity increase but the Hurst exponent decreases with increasing gas velocity. Whereas, the bubble chord length decreases, but the Hurst exponent increases with an increase in liquid velocity in the continuous bubble column (U_L>0.02 m/s). The Hurst exponent has been found to have a definite relationship with the bubble chord length and its standard deviation.     

Local bubble behavior in three-phase fluidized beds

Bubble behavior, including bubble Sauter diameter, bubble rise velocity, bubble frequency and local gas holdup in different radial and axial positions, was measured using a dual electro-conductivity probe in air-water-glass beads fluidization systems. It has been found that the bubble characteristics differ significantly in various flow regimes, depending on the operating conditions; the radial distribution of bubble parameters also changes from one flow regime to another. Thus, it is necessary to employ local bubble behavior in the modeling of three-phase fluidized beds.     

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.     

HEAT TRANSFER CHARACTERISTICS OF THREE PHASE FLUIDIZED BEDS

Heat transfer characteristics of two (liquid-gas, liquid-solid) and three (liquid-gas-solid) phase fluidized beds have been studied in a 15.2 cm-ID column fitted with an axially mounted cylindrical healer. Effects of gas velocity (0-12 cm/s). liquid velocity (0-16cm/s), particle size (1.7-8.0 mm) and liquid viscosity (0.001-0.039 Pa s) on heat transfer coefficient were determined. The heat transfer coefficient increased with fluid velocities and particle size and it decreased with liquid viscosity in two and three phase fluidized beds. The bed porosity at which the maximum heat transfer occurred decreased with particle size but increased with liquid viscosity. The coefficient were correlated in terms of experimental variables. Modified Nusselt number from the present and previous studies has been correlated with modified Prandtl and Reynolds numbers.