Hydrodynamics of a high pressure three‐phase fluidized bed subject to foaming

The effects of pressure and surfactants on the phase holdups and flow regime transition velocities of gas–liquid–solid fluidized beds were investigated. The effect of pressure on the bed phase holdups is significant and more pronounced at larger gas flow rates where pressure has a greater effect on the equilibrium bubble size. The addition of a surfactant leads to an increase in the gas holdup and a lowering of the solids and liquid holdups. The presence of a surfactant with a liquid flow results in shearing of the bubbles across the gas–liquid distributor, limiting the effect of pressure. Finally, for all conditions, gas holdups in the freeboard region were greater than in the bed.     

Hydrodynamics of a gas-liquid-solid three phase circulating fluidized bed

The hydrodynamics of a gas-liquid-solid circulating fluidized bed was investigated. A new regime, the three phase circulating fluidization regime, was discovered for the first time. The characteristics of this regime were compared with that of the conventional fluidization regime and the transport regime. The particle circulation rate and the gas and solids holdups in circulating fluidization regime were studied.     

Minimum fluidization velocity and gas holdup in gas–liquid–solid fluidized bed reactors

Experiments were performed to study the hydrodynamics of a cocurrent three‐phase fluidized bed with liquid as continuous phase. Based on the 209 experimental data (with four liquid systems and five different particles) along with 115 literature data from six different sources on minimum fluidization velocity, a unique correlation for the estimation of minimum fluidization velocity in two‐phase (u_g = 0) as well as in three‐phase systems is developed. A data bank consisting of 1420 experimental measurements for the fractional gas phase holdup data with a wide range of variables is used for developing empirical correlations. Separate correlations are developed for two flow regimes observed in this present work. The proposed correlations are more accurate and simpler to use. © 2002 Society of Chemical Industry     

Fluidization regimes in two‐ and three‐phase fluidized beds: Comparison between measurement techniques

Fluidization regimes were studied in liquid‐solid and gas‐Liquid‐solid fluidized beds. The liquid velocities at which regime transitions occur in liquid‐solid and gas‐liquid‐solid systems were obtained for monosize and multicomponent systems. Minimum fluidization, complete fluidization and complete mixing velocities of particles were obtained from pressure drop measurements, a collision technique or a conductivity method. The collision technique provided accurate complete fluidization and complete mixing velocities by measuring directly the particle motion. The conductivity technique, which is easier to implement, gave similar but less accurate results.     

Methanol synthesis in a three‐phase catalytic bed under nonwetted condition

To overcome the heat removal problem encountered in methanol synthesis at high syngas concentrations in the gas phase, a three‐phase nonwetted catalytic system was established by introducing an inert liquid medium into a fixed‐bed reactor. To form a repellent interface between the liquid and the catalyst, the catalyst was modified into hydrophobic, while the liquid medium was chosen as a room temperature ionic liquid with hydroxyl groups. The liquid‐solid contact angle was measured to be 115°, and only 20% of the catalyst external surface was wetted by the liquid. Under three‐phase condition, the reaction rate was measured to be 60%–70% of gas‐phase reaction, while it was merely 10%–20% for the fully wetted catalyst. From the resistance analysis on the mass transfer and reaction steps, the overall reaction rate is expected to increase further if the surface could be more wet proofed. © 2016 American Institute of Chemical Engineers AIChE J, 63: 226–237, 2017     

Effects of oxygen supply condition and specific biofilm interfacial area on phenol removal rate in a three‐phase fluidized bed bioreactor

The effects of oxygen supply conditions and specific biofilm interfacial area on the phenol removal rate in a three‐phase fluidized bed bioreactor were evaluated. The experimental data were well‐explained by the semi‐theoretical equation based on the assumption that the reaction rate follows first‐order reaction kinetics with respect to oxygen and zero‐order one with respect to phenol. Two cases, biological reaction as rate‐controlling step and oxygen absorption as rate‐controlling step, were both explicable by this semi‐theoretical equation. The maximum volumetric phenol removal rate was 27.4 kg·m^?3·d^?1.     

Three‐phase fluidized distillation

Separation efficiency in distillation operations can be improved by modifying the characteristics of the dispersions formed on the trays. The present work reports on the hydrodynamic and mass transfer characteristics of liquid‐solid‐vapour dispersions formed on sieve trays without downcomers of a distillation column operating under total reflux conditions. Murphree efficiency and the operating limits of distillation using the ethanol‐n‐butanol system are analyzed for a large range of vapour velocities and liquid mixture compositions, utilising wettable PVC particles and non‐wettable silicone, PE, and Teflonr? particles. It was verified that wettable particles show a drastic reduction in the upper operating limit of vapour velocity, but this does not occur for non‐wettable particles. Tray efficiency can be increased when non‐wettable particles are used, mainly for high vapour velocity operations.     

Bubble behavior in corrugated‐wall bubbling fluidized beds—Experiments and CFD simulations

A new concept to harness bubble dynamics in bubbling fluidization of Geldart D particles was proposed. Various geometrical declinations of a cold‐prototype corrugated‐wall bubbling fluidized bed were compared at different flow rates (U_g) to conventional flat‐wall fluidized bed using high‐speed digital image analysis. Hydrodynamic studies were carried out to appraise the effect of triangular‐shaped wall corrugation on incipient fluidization, bubble coalescence (size and frequency), bubble rise velocity, and pressure drop. Bubble size and rise velocity in corrugated‐wall beds were appreciably lower, at given U_g/U_mb, than in flat‐wall beds with equal flow cross‐sectional areas and initial bed heights. The decrease (increase) in size (frequency) of bubbles during their rise was sustained by their periodic breakups while protruding through the necks between corrugated plates. Euler‐Euler transient full three‐dimensional computational fluid dynamic simulations helped shape an understanding of the impact of corrugation geometry on lowering the minimum bubbling fluidization and improving gas distribution. © 2011 American Institute of Chemical Engineers AIChE J, 2012     

Experimental profiles of lateral mixing of feed particles in a three‐dimensional fluidized bed

Solids mixing data of high quality is one of the most crucial steps for quantitative studies, but it is a difficult task to obtain in a fluidized bed especially with a 3D configuration. Therefore a novel sampling technique is developed with bed collapse method, for measuring lateral mixing of feed particles in a 3D fluidized bed. The sampling tool is designed using a “bottom‐to‐top sampling” idea. Its development, configuration and measurement repetition are discussed in detail. The effects of mixing time, fluidizing gas velocity, and particle size of bed material on the tracer distribution are investigated. A quantitative comparison of lateral dispersion coefficient shows that our results agree fairly well with measurements and predictions of correlations for lab‐scale fluidized systems in previous studies. The presented 2D profiles of the lateral mixing can be used to validate fundamental solids mixing models or verifying convenient measurement techniques. © 2010 American Institute of Chemical Engineers AIChE J, 2011     

Dispersed phase characteristics in three phase (liquid-liquid-solid) fluidized beds

Two ideal droplet length (l,) distributions have been derived for two different droplet shapes. The dispersed phase holdup (?_d) increases with increasing dispersed phase velocity (U_d), but decreases with increasing continuous phase velocity, (U_c) in three-phase fluidized beds. In the droplet-coalescing flow regime, l_v and the droplet rising velocity (v_d) increase, but the spherical droplet fraction (k) decreases with increasing U_d and u_c. In the droplet-disintegrating flow regime, the effects of u_d and U_c on l_v and k are insignificant, but v_d increases with increasing U_c. Maximum values of l_v, occur in the bed containing 1.7 mm diameter particles and l_v has an uniform length of around 2.0 mm in beds with particle size larger than 3.0 mm.     

Axial distribution of solids holdup for both hydrophilic and hydrophobic particles in three‐phase fluidized beds

The axial concentration distribution of the hydrophilic and hydrophobic particles, corresponding to poor and good attachment to rising bubbles in water, was investigated in a gas‐liquid‐solid fluidized bed. The sedimentation‐dispersion model depicted satisfactorily the axial distribution of solids holdup by modifying only the terminal settling velocity, which considered the effect of an additional drag force resulting from the attached rising bubbles. This model explained correctly the different phenomena, i.e. solids holdup decreased with increasing the axial height in the non‐attached system, but increased in the attached system.     

New two‐phase and three‐phase Rachford‐Rice algorithms based on free‐water assumption

Based on the free‐water assumption that the water‐rich liquid phase contains only pure water, we develop two simple free‐water Rachford‐Rice methods: 1) a two‐phase free‐water Rachford‐Rice method where the phase fractions can be analytically solved; and 2) a three‐phase free‐water Rachford‐Rice method where there is only one unknown in the objective function (i.e. the vapour‐phase fraction) that is used for solving the phase fractions. Combining these two Rachford‐Rice methods, a new free‐water flash algorithm is developed to perform multiphase flash calculations where single‐phase equilibria, two‐phase equilibria, and three‐phase vapour‐liquid‐aqueous equilibria can be considered. In this free‐water flash algorithm, we first test if the mixture is stable; if the mixture is found to be unstable, we directly initiate the three‐phase free‐water flash. A set of criteria is developed for one to properly switch from a three‐phase free‐water flash to either a two‐phase free‐water flash or a conventional two‐phase flash, depending on whether a water‐rich phase is present in the two‐phase equilibrium. We also develop efficient and robust methods for initializing the equilibrium ratios for the two‐phase flashes. The negative flash is allowed in the flash calculation algorithms. A number of example calculations are carried out to demonstrate the robustness of the newly developed algorithm. A good agreement can be achieved between the flash results obtained by the new flash algorithm and those obtained by the conventional full three‐phase flash algorithm.     

Scale‐dependent nonequilibrium features in a bubbling fluidized bed

We investigate experimentally the nonequilibrium features in a pseudo 2‐D bubbling fluidized bed. Velocities of individual particles are measured by using a particle tracking velocimetry (PTV) method, and void fractions are obtained with the Voronoi tessellation. A bimodal shape of probability density function (PDF) for particle vertical velocity is found in not only time‐averaged but also time‐varying statistics, which is caused by the transition between the dense and dilute phases and breaks the local‐equilibrium assumption in continuum modeling of fluidized beds. The results of time‐varying radial distribution function and voidage distribution also confirm this finding. Moreover, the analysis of voidage, particle velocity, granular temperature and turbulent kinetic energy of particles shows that there is no scale‐independent plateau over the interface, and it seems hard to find a scale‐independent plateau to separate the micro‐ and meso‐scales of fluidized beds, which require sub‐grid meso‐scale modeling for continuum or coarse‐graining methods of gas‐fluidized systems. © 2018 American Institute of Chemical Engineers AIChE J, 64: 2364–2378, 2018     

Bubble splitting under gas–liquid–liquid three‐phase flow in a double T‐junction microchannel

Gas–aqueous liquid–oil three‐phase flow was generated in a microchannel with a double T‐junction. Under the squeezing of the dispersed aqueous phase at the second T‐junction (T2), the splitting of bubbles generated from the first T‐junction (T1) was investigated. During the bubble splitting process, the upstream gas–oil two‐phase flow and the aqueous phase flow at T2 fluctuate in opposite phases, resulting in either independent or synchronous relationship between the instantaneous downstream and upstream bubble velocities depending on the operating conditions. Compared with two‐phase flow, the modified capillary number and the ratio of the upstream velocity to the aqueous phase velocity were introduced to predict the bubble breakup time. The critical bubble breakup length and size laws of daughter bubbles/slugs were thereby proposed. These results provide an important guideline for designing microchannel structures for a precise manipulation of gas–liquid–liquid three‐phase flow which finds potential applications among others in chemical synthesis. © 2017 American Institute of Chemical Engineers AIChE J, 63: 376–388, 2018     

Suspension of solid particles in multi‐impeller three‐phase stirred tank reactors

Solids suspension characteristics in gas—liquid–solid three‐phase stirred tanks with multi‐impellers were experimentally examined. Minimum impeller speeds for ultimately homogeneous solid suspension have been measured stirred tank reactors. Three impellers were installed: two four‐pitched blade downflow disk turbines and one Pfaudler type impeller chosen to provide good gas dispersion and to accomplish off‐bottom suspension of solid particles, respectively. Gas dispersion causes an increase in particle sedimentation associated with a decrease in power consumption and as a result, minimum impeller speeds for ultimately homogeneous solid suspension increase with increasing gas flow rates. A correlation was developed to predict minimum impeller speeds for ultimately homogeneous solid suspension. The proposed correlation, which agrees satisfactorily with the experimental results, is expected to be useful in design and scale‐up.     

Transition velocities and phase holdups at minimum fluidization in gas–liquid–solid systems

Hydrodynamic experiments were performed using a 127‐mm diameter column with 3.2‐mm porous alumina, 3.3‐mm polymer blend, 5.5‐mm polystyrene and 6.0‐mm glass spheres, with water, aqueous glycerol solution and silicone oil as liquids, and air as the gas. The voidage at minimum fluidization fell initially to a minimum, then rose gradually with increasing superficial gas velocity, and was lower for three‐phase systems than for corresponding two‐phase (liquid–solid) fluidized beds. The compaction appears to be due to agitation by gas bubbles near the minimum liquid fluidization condition. The gas holdups agree reasonably well with the correlation of Yang et al. (1993). Curves of minimum liquid fluidization velocity, U_lmf, vs. superficial gas velocity, U_g always show U_lmf decreasing as U_g increases, initially in a concave‐downward manner, but sometimes concave‐upward.     

PDA measurements in a three‐phase bubble column

Phase Doppler anemometry was used to quantify the flow characteristic of a three phases (liquid, solid, and bubbles) cylindrical bubble column driven by a point air source made of a 30‐mm diameter perforated air stone centrally mounted at the bottom. The cylindrical bubble column had an inner diameter of 152 mm and was filled with liquid up to 1 m above the point source. Acrylic beads with a nominal diameter of 3 mm were used as the solid phase. To match the density of the solid phase which was 1.05 kg/m³, the liquid density was raised to about 1.0485 kg/m3 by added salt. The bubble diameters generated were within the range of 600–2400 µm. The detailed turbulent characteristics of the liquid‐phase velocity, bubble diameter, bubble velocity, and solid velocity were measured at three different air rates, namely 0.4, 0.8, and 1.2 L/min (corresponding to average gas volume fraction of 0.0084, 0.0168, and 0.0258, respectively) for the homogeneous bubble column regime. With the addition of the solid phase, the flow field was found to be relatively steady compared to the two‐phase column referencing the probability density functions for both the liquid and bubble velocities. An analysis based on the determination of the drag forces and transversal lift forces was performed to examine the flow stability in the three‐phase bubble column. The analysis illustrated that how the added solid phase effectively stabilized the flow field to achieve a steady circulation in the bubble column and a generalized criterion for the flow stability in the three‐phase bubble column was derived. Further investigation for the transition and the heterogeneous bubble column regime with air rates at 2.0 and 4.0 L/min shown that this criterion can also be used as a general prediction of flow stability in this three‐phase bubble column. © 2013 American Institute of Chemical Engineers AIChE J, 59: 2286–2307, 2013     

Characterizing bubble behavior in a pseudo-2D fluidized bed of wet particles

By applying digital image analysis on the bubble characteristics in a two dimensional wet-particle fluidized-bed, we report two-stage evolution of bubble characteristics with increasing liquid content. In the first stage, bubble number and uniformity of bubble fraction increase, while bubble average diameter and aspect ratio decrease. In the second stage, these characteristics shift toward an opposite direction. This two-stage evolution of bubble characteristics are analogous to that of reducing particle size in dry-particle fluidization, and the fluidizing properties of particles shifts from Geldart Group B to Group A and then to Group C. Furthermore, liquid addition causes a continuous decrease of bubble fraction and bubble flow rate. This is different from dry-particle fluidization, in which reducing particle size causes an increase trend. An explanation for this difference is that liquid addition increases the equivalent size of agglomerates in wet-particle fluidization, which is opposite to the effect of reducing particle size.     

Modified gas‐perturbed liquid model (GPLM) for minimum fluidization velocity in a two‐phase inverse fluidized bed reactor with Newtonian and non‐Newtonian systems

In the present investigation minimum fluidization velocity, U_mf, in a two‐phase inverse fluidized bed reactor is determined using low‐density polyethylene and polypropylene particles of different diameters (4,6 and 8 mm) by measuring pressure drop. In a glycerol system U_mf decreased gradually with increase in viscosity up to a value of 6.11 mPa s (60%) and on further increase there was a slight increase in U_mf. In the case of the glycerol system the U_mf was found to be higher when compared to water. In the non‐Newtonian system (carboxymethylcellulose), U_mf decreased with increase in concentration in the range of the present study. The U_mf was found to be lower when compared to water as liquid phase. The modified gas‐perturbed liquid model was used to predict the minimum fluidization liquid velocity (U_lmf) for Newtonian and non‐Newtonian systems. Copyright © 2006 Society of Chemical Industry