Hydrodynamics and mass transport in rotary disk contactors

This paper reports a detailed study of the hydrodynamics involved in the operation of rotary disk contactors (RDC). New information is presented regarding the ‘critical rotor speed’ which divides the two hydrodynamic regimes of near constant and varying characteristic velocity of the dispersed droplet phase. Generalized correlations for prediction of characteristic velocity both above and below the critical rotor speed have been proposed under conditions of no solute transfer as well as solute transfer between the dispersed and continuous phases. Experimental data on mass transfer obtained in the two hydrodynamic regimes of RDC operation have also been correlated by the application of mass transport theory taking into account the factors affecting the interfacial area of contact, drop size, characteristic velocity and dispersed phase holdup.     

Design and characterization of a multistage,mechanically stirred column absorber

A multistage, mechanically stirred column absorber has been designed and built with a modular construction, based on preliminary experiments with a test column. The column has been characterized as a gas-liquid contactor by its gas holdup, gas and liquid axial dispersion, mixing times, oxygen transfer coefficients and power consumptions, determined as a function of gas velocity, liquid velocity and impeller speed for one and two impellers per stage.Gassed power was correlated with ungassed power, gas rate and impeller speed. The gas phase axial mixing was essentially plug flow and the liquid phase axial mixing varied between 5 and 12 equivalent stages.Oxygen transfer coefficients were correlated with power consumptions and aeration rates by the equation K_La γ (P/V)^a(υ_sg)^b. The oxygen transfer coefficients with single stiffer stages were 25% above those for the double stirrer stages for equal power consumption and gas rates. Except for the low aeration and high power consumption extremes, the column showed superior oxygen transfer performance. in comparison to tubular loop and tank fermenters.     

Radial gas mixing characteristics in a downer reactor

In a downer reactor (0.1 m-I.D.x3.5 m-high), the effects of gas velocity (1.6-4.5 m/s), solids circulation rate (0–40kg/m²s) and particle size (84, 164 Μm) on the gas mixing coefficient have been determined. The radial dispersion coefficient(D _r ) decreases and the radial Peclet number (Pe_r) increases as gas velocity increases. At lower gas velocities, D_r in the bed of particles is lower than that of gas flow only, but the reverse trend is observed at higher gas velocities. Gas mixing in the reactor of smaller particle size varies significantly with gas velocity, whereas gas mixing varies smoothly in the reactor of larger particle size. At lower gas velocities, D_r increases with increasing solids circulation rate (G_s), however, D_r decreases with increasing G_s at higher gas velocities. Based on the obtained D_r values, the downer reactor is found to be a good gas-solids contacting reactor having good radial gas mixing.     

Radial gas mixing in a fluidized bed using response surface methodology

Radial gas mixing in a fluidized bed was studied using response surface methodology (RSM), which enables effect examinations of parameters with a moderate number of experiments. All experiments were conducted in a 0.29-m ID fluidized-bed cold model. The gas dispersion process within the bed is described using the dispersed plug flow model. Pure carbon dioxide was used as the tracer gas, continuously injected into the center of the bed by a point source. The downstream radial tracer concentration profile was measured using a gas chromatograph.The radial gas dispersion coefficient, D_r, was well correlated with operating parameters and the particle and gas properties: (U−U_mf)/U_mf, H_s/d_b, φ_d, and Ar, with a determination coefficient R² of 0.966. Effect test indicates that the dimensionless characteristic velocity, (U−U_mf)/U_mf, has the most significant influence on D_r, while the static bed height to bed diameter ratio, H_s/d_b, is less remarkable. The interactions of (U−U_mf)/U_mf with the distributor open-area ratio, φ_d, and with the Archimedes number, Ar, both play important roles. An evolutive response surface model was proposed to describe the radial gas mixing in the bubbling/slugging fluidization regimes.     

提升管与气-固环流床层耦合反应器颗粒相循环比的模型及预测

引言Geldart A类颗粒气-固环流技术是一种利用气-液环流原理,并结合气-固聚式流化体系特点而开发的一种新型流态化技术[1],具有气固接触效率高、传质传热性能好等优点,被广泛用于石油炼制领域中的催化裂化汽提器、提升管出口粗旋快分的预汽提器、外取热器、石油焦燃烧器和降烯烃反应器等装备中[2]。     

A study on the flow characteristics in a pulsed doughnut-disc type plate extraction column

The axial dispersion coefficients in the continuous phase and holdup of dispersed phase have been studied in a 4.2 cm inside diameter and 200 cm height pulsed doughnut-disc type plates extraction column. The axial concentration gradient in a continuous extraction column was expressed mathematically in terms of Peclet number by axial dispersion model. Peclet numbers have been calculated from response curves using KC1 solution as an impulse input fracer. Experimental data have been taken for both continuous and dispersed phase with plate spacing, pulsing amplitudes, frequencies, and superficial velocities as system variables. Modified axial dispersion coefficients have been correlated by regression analysis of experimental data, and following equations were obtained. 1. Axial dispersion coefficient (single phase) E_c = 3.5H^-13 A^1.5.1 f + 30.95 U_c 2. Axial dispsion coefficient (two phase) E_c = 2.36 H^{-0 8} A^1.34 f + 20.89 U_c 3. Fractional holdup of the dispersed phase Φ_d = 4 2xl0^-5H^-0.44 Af^1.28U_d ^0.93     

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).     

Solids mixing in a down-flow circulating fluidized bed of 0.418-m in diameter

The axial and lateral solids mixing in a down-flow circulating fluidized bed of 0.418-m diameter was investigated by a pneumatic injection phosphor tracer technique (PIPTT). The axial and lateral solids dispersion were determined by measuring the solids RTD at same axial but different lateral positions using point sources for tracer injection. A two-dimensional dispersion model described the measured RTD curves satisfactorily. The results were compared to those obtained in the small scale downers and the scale-up effect was investigated. The axial solids Peclet number Pe_a is around 110 and invariable with changing U_g, G_s and ?_s, while the lateral solids Peclet number Pe_r is linearly increasing with ?_s. And Pe_r is found to decrease with the square root of inner diameter (ID) in comparison with the results obtained in small ID downers. Correlation of Pe_r, Pe_r = (15 + 70.7 ?_s)D^− 0.5, is proposed.     

Critical gas velocity in a three‐phase internal loop airlift reactor with low‐density particles

In the present investigation the effects of the addition of organic additives (propanol, benzoic acid, iso‐amyl alcohol and carboxymethyl cellulose) on the critical gas velocity, (U_sg)_c, in an internal airlift loop reactor with low‐density particles (Nylon‐6 and polystyrene) were reported. Whereas the (U_sg)_c was reduced by adding the above additives, it increased with solids loading and density of the particles. The draft tube‐to‐reactor diameter ratio (D_E/D) in the range of 0.5–0.6 gave minimum (U_sg)_c values. The proposed dimensionless correlation predicted the experimental data well. Copyright © 2004 Society of Chemical Industry     

CFD modeling of flow,macro-mixing and axial dispersion in a bubble column

The flow pattern in a bubble column depends upon the column diameter, height, sparger design, superficial gas velocity and the nature of gas–liquid system. In this paper, the effect of some of these parameters have been simulated using Computational Fluid Dynamics (CFD). The relationship of these parameters with the interphase force terms has been discussed. A complete energy balance has been established. Using this methodology, the flow patterns reported by Hills (1974), Menzel et al. (1990), Yao et al. (1991) and Yu and Kim (1991) have been simulated. Excellent agreement has been shown between the CFD predictions and the experimental observations. The above model has been extended to homogenization of an inert tracer. In order to confirm this model, mixing experiments were carried out in a 200 mm i.d. bubble column. A radioactive tracer technique was used for the measurement of mixing time. Tc-99m (^99m Tc), in the form of sodium pertechnate salt, was used as the liquid phase tracer. Good agreement has been shown between the predicted and the experimental values of mixing time. The model was further extended for the estimation of axial dispersion coefficient (D_L). Excellent agreement between the simulated and the experimental values of the axial dispersion coefficient confirms the predictive capability of the CFD simulations for the mixing process.     

Effect of system properties on the performance of liquid-liquid extraction columns—1: Packed column

Mean drop size, fractional hold-up of dispersed phase, and axial mixing have been determined in a 72 mm dia. packed column, packed with 8 mm glass Raschig rings, using the systems toluene-acetone-water and MIBK-acetic acid-water, in an attempt to assess the effects of the solute transfer processes on column operation.For the MIBK-system, where the packing size corresponded to the normal recommended conditions d_p > d_FC (packing size greater than the critical size of packing), solute presence and the direction of solute transfer was found to affect drop size, fractional hold-up and the axial mixing in the dispersed phase appreciably, in a manner which was consistent with either suppressed or enhanced coalescence characteristics according to the appropriate changes in solute system.For the toluene system, however, with the packing size corresponding to the condition d_p < d_FC, the drop behaviour was dominated by the action of the packing void spaces and mean drop size was independent of solute presence and direction of transfer.In both cases, axial mixing in the continuous phase was independent of solute effects, except in so far as these changes modified hold-up of the dispersed phase.     

Mass transfer to a drop stream from a field liquid

Mass transfer from a stream of drops falling freely in a stagnant liquid was investigated. Drop streams were produced by a dripping method and by a jet breakup method. Water and isobutanol, mutually saturated, were used as the dispersed and the continuous phases. Sodium hydroxide was transferred from isobutanol to water drops which were initially free of solute. The mass transfer resistance is on the continuous phase side. The mass transfer coefficient and terminal velocity of drop streams were measured experimentally. The experimental results show that the mass transfer coefficient in the drop stream is affected by the shielding effect of the previous drops. The experimental data have been correlated as K_t/U_t^0.5 versus interdrop distance l, a relationship describing the effect of the interdrop distance on the mass transfer coefficient in the continuous phase.     

Fluidization of Group B particles with a rotating distributor

A novel rotating distributor fluidized bed is presented. The distributor is a rotating perforated plate, with 1% open-area ratio. This work evaluates the performance of this new design, considering pressure drop, Δp, and quality of fluidization. Bed fluidization was easily achieved with the proposed device, improving the solid mixing and the quality of fluidization.In order to examine the effect of the rotational speed of the distributor plate on the hydrodynamic behavior of the bed, minimum fluidization velocity, U_mf, and pressure fluctuations were analyzed. Experiments were conducted in the bubbling free regime in a 0.19 m i.d. fluidized bed, operating with Group B particles according to Geldart's classification. The pressure drop across the bed and the standard deviation of pressure fluctuations, σ_p, were used to find the minimum fluidization velocity, U_mf. A decrease in U_mf is observed when the rotational speed increases and a rise in the measured pressure drop was also found. Frequency analysis of pressure fluctuations shows that fluidization can be controlled by the adjustable rotational speed, at several excess gas velocities.Measurements with several initial static bed heights were taken, in order to analyze the influence of the initial bed mass inventory, over the effect of the distributor rotation on the bed hydrodynamics.     

双转子连续混炼机混合工艺特点及影响因素

讨论了双转子连续混炼机的混合工艺特点及转子结构尺寸，转子螺杆顶部线速度，物料卸料门开启度大小对混合过程的影响，提出了以物料在混炼机中的转向混合循环次数，转子螺棱顶部线速度为基本参数控制混炼机的混合过程。     

HYDRODYNAMICS IN YORK-SCHEIBEL COLUMNS

Dispersed phase hold-up, hold-up profile, drop-size distribution and phase inversion were investigated in a third type of Scheibel column (7.6 cm diameter, 18 stage) with toluene-water as the liquid-liquid system. The concept of characteristic drop velocity was used to describe the relationship between dynamic hold-up and two phase flow rates. The axial dispersed phase hold-up and the mean drop size varied significantly with rotor speed and column height. In addition, the characteristics of phase inversion was explained by Sarkar's model.     

Prediction of the volumetric mass transfer coefficient in pneumatic contactors

Semi-theoretical expressions are developed for the prediction of the volumetric mass transfer coefficient, (K_La_D)_T, in pneumatic contactors, using the correlation of Calderbank and Moo-Young Chem. Engng Sci.16, 39 (1961) for the mass transfer coefficient and the local isotropic turbulence theory for predicting bubble diameter. A direct proportionality of (K_La_D)_T to the gas hold-up is predicted, with an exponent of 1.2 on ε_G, while, in terms of physical parameters, (K_La_D)_T is predicted to be proportional to the 0.8 power of the superficial gas velocity for both bubble column and airlift contactors, and proportional to (1 + A_d/A_r)⁻² for airlift contactors, where A_d/A_r is the downcomer-to-riser cross-sectional area ratio. Experimental results obtained in bubble column and airlift contactors (external-loop and concentric-tube) of pilot-plant scale (ca. 50L liquid capacity), with water and 0.15 kmol m⁻³ NaCl solution as liquid media, were used to test the proposed expressions.     

Dependence of particle fluctuation velocity on gas flow, and particle diameter in gas fluidized beds for monodispersed spheres in the Geldart B and A fluidization regimes

In this paper we present new experimental data on the steady-state, mean squared, fluctuation velocity, or granular temperature, of Geldart B polymer, glass, nickel, and stainless steel monodispersed spheres averaged over the wall of a gas fluidized bed, as a function of gas flow and sphere diameter. The granular temperature is obtained by Acoustic Shot Noise technology—namely power spectral analysis of the steady state vibrational energy of the wall excited by random sphere impact, and calibrated by hammer excitation over the wall. The new data extends to polymer and metallic spheres the experimental discovery of a 1996 paper of Cody et al. that the fluctuation velocity of Geldart B glass spheres when scaled to the gas superficial velocity, U_s, is inversely proportional to sphere diameter, directly proportional to a fundamental length scale, D_o^B, and is a universal function of U = (U_s / U_mf). We also demonstrate that the new data is consistent with the diameter dependence of the fluctuation velocity that can be derived from both the 1997 paper of Menon and Durian, who measured random sphere motion near the wall through the spectroscopy of scattered laser light, and the 1992 paper of Rahman and Campbell, who measured the average granular pressure of random sphere impact on a porous steel membrane. While the inverse scaling of the fluctuation velocity with sphere diameter, and the existence of a fundamental length scale for gas fluidization, D_o^B, had not been a feature of any published fundamental model, or computer simulation, of the steady state granular temperature of spheres in gas fluidized beds, we show that it is a feature of two recent dense kinetic fluidization models published in 1999, by Buyevich and Kapbasov, and Koch and Sangani. Both theories implicitly define a fundamental length scale for the fluctuation velocity, D^? = (μ_f² / ρ_p²g)^1 / 3, where ρ_p is the sphere density, μ_f is the gas viscosity, and g is the laboratory gravitational field. The new data for polymer, glass, nickel and stainless steel spheres presented in this paper, defines D_o^B = (56 ± 2)D^?. We use the Anderson-Jackson stability model to show that the length scale D_o^B, also defines a stability length scale, such that for D < D_o^B(D > D_o^B), the uniform dense phase of the fluidized bed is stable (unstable), against one dimensional, first order fluctuations in sphere concentration. The length scale, D_o^B is thus the theoretical equivalent to the empirical scaling length introduced by Geldart, D_B/A, to distinguish spheres (D > D_B/A) that bubble at fluidization, from spheres (D < D_B/A) that fluidize before bubbling. Finally, we present new experimental data, on the remarkable changes in the granular temperature, bed expansion, and bed collapse time, between Geldart B and Geldart A monodispersed glass spheres, and compare that data to granular temperature, and bed expansion, for Geldart A rough, non-spherical, log-normal dispersed diameter catalytic particles.     

An oscillating baffle column fitted with stator rings III. Hold-up and drop distribution

A study of drop sizes and their distribution in a new extractor shows the distribution to obey a log-normal law. Hold-up of dispersed phase obtained manometrically is related to oscillation speed (N), relative flow-rate (Q_t) and phase flow ratio. Equations are proposed for prediction of d₃₂ and Ψ from N and Q_t. The incipient flood point can be established either from the graphs or equations. An equation proposed by Misek to correlate the relative flow-rate of the phases, the interdroplet collision rate, and coalescence has been examined. Values of U₀, the free fall velocity obtained indirectly by experiment from the Misek-type equation, are compared with free fall velocities for single drops obtained by other workers. The system examined was water-acetic acid (solute)-carbon tetrachloride (disperse phase).     

Optical measurements of porosity and fluid motion in packed beds

A marker tracking method is used to measure the radial distributions of porosity and fluid motion in transparent packed beds of equilateral cylinders. Data are given for particle Reynolds numbers from 5 to 280 in beds with D/D_p = 10.7 and L/D_p = 20.6. Spatial oscillations with period ∼ D_p are observed for the smoothed porosity ε(r), interstitial axial velocity u(r), and local superficial velocity u(rε(r). The latter velocity attains its global maximum and minimum at distances near 0.2D_p and 0.5D_p from the wall. The frequency distributions of the marker displacements are also analysed and compared with a smoothed diffusion model.