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.     

Transition to turbulent fluidization in a binary solids fluidized bed

This paper presents a study on the transition velocity from bubbling to turbulent fluidization in a binary solids fluidized bed. Experiments were carried out with two kinds of binary solids mixtures with FCC as fine particles and silica sands as coarse particles. The onset velocity to turbulent fluidization, U_c, determined by the measurement of pressure fluctuations, was found to increase with increasing the fraction of coarse/heavy solids. By introducing an equivalent particle diameter and an equivalent particle density, the results obtained in this study can properly be described by a general correlation of U_c proposed by Cai and co-workers (1989) for mono-density particles with relatively narrow size distribution.     

Dynamics of segregation and fluidization of binary mixtures in a cylindrical fluidized bed

Dynamics of segregation and fluidization of unary particles and binary mixtures in a cylindrical fluidized bed is investigated using temporally– and spatially–resolved measurements of solids volume fraction (α_s) performed using Electrical Capacitance Tomography (ECT). Through the comparison with high-speed imaging, we have shown that ECT can be used to measure the segregation behavior in cylindrical fluidized beds quantitatively. ECT measurements have been used further to quantify the effects of mixture composition, particle–diameter ratio, and superficial gas velocity on the bed segregation behavior. Dynamics of fluidization behavior is characterized using the time–evolution of local α_s fluctuations, corresponding frequency distribution, and bubble size distribution. Further, a relation between the measured variance of α_s fluctuations at different radial locations and corresponding flow structures under different fluidization conditions is established. The present work helps to understand dynamics of segregation and fluidization of binary mixtures and to provide a database for validation of Eulerian multifluid CFD models.     

Minimum fluidization velocity in gas‐liquid‐solid minifluidized beds

The initial fluidization characteristics of gas‐liquid‐solid minifluidized beds (MFBs) were experimentally investigated based on the analyses of bed pressure drop and visual observations. The results show that U_Lmf in 3–5 mm MFBs can not be determined due to the extensive pressure drop fluctuations resulting from complex bubble behavior. For 8–10 mm MFBs, U_Lmf can be confirmed from both datum analyses of pressure drop and Hurst exponent at low superficial gas velocity. But at high superficial gas velocity, U_Lmf was not obtained because the turning point at which the flow regime changes from the packed bed to the fluidized bed disappeared, and the bed was in a half fluidization state. Complex bubble growth behavior resulting from the effect of properties of gas‐liquid mixture and bed walls plays an important role in the fluidization of solid particles and leads to the reduction of U_Lmf. An empirical correlation was suggested to predict U_Lmf in MFBs. © 2016 American Institute of Chemical Engineers AIChE J, 62: 1940–1957, 2016     

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     

Flow structure and thickness of annular downflow layer in a circulating fluidized bed riser

Flow structures were determined in a circulating fluidized bed (CFB) riser (0.203 m i.d.×5.9 m high) of FCC particles (d_p=70 μm, ρ_s=1700 kg/m³). A momentum probe was used to measure radial momentum flux profiles at several levels and to distinguish between upward and downward flow regions. Time-mean dynamic pressure (ΔP_m) decreases towards the wall in the range U_g=5-8 m/s, G_s=10-340 kg/m² s. The thickness of the annular downflow layer based on ΔP_m=0 reaches a maximum with increasing height. The annular downflow layer disappears locally with increasing solids mass flux (G_s) at a constant gas velocity, with achievement of the dense suspension upflow (DSU) regime. A new correlation is developed to predict the time-mean thickness of solids down-flowing layer based on solids mass flux and momentum flux. It successfully accounts for the variation of the annular layer thickness with height and G_s, and covers a wide G_s range right up to near the onset of the DSU regime.     

Experimental Study on Bubble Rising and Descending Velocity Distribution in a Slurry Bubble Column Reactor

To determine bubble rising and descending velocity simultaneously, a BVW‐2 four‐channel conductivity probe bubble parameters apparatus and its analysis are used in gas‐liquid and gas‐liquid‐solid bubble columns. The column is 100 mm in internal diameter and 1500 mm in height. The solid particles used are glass beads with an average diameter of 17.82 μm, representing typical particle size for catalytic slurry reactors. The effects of superficial gas velocity (1.0 cm/s ≤ U_g ≤ 6.4 cm/s), solid holdup (0 % ≤ ?_s ≤ 30 %), and radial location (r/R = 0, 0.4, and 0.7) on bubble velocity distributions are determined. It is found that increasing U_g can increase the velocity of bubbles but do not exert much influence on bubble velocity distribution. Solid holdup mainly affects the distribution of bubble velocity while the radial direction affects bubble velocity distribution only slightly. The ratio of descending bubbles to rising bubbles increases from the bubble column center to the wall. It can be proved experimentally that large bubbles do not always rise faster than small bubbles at higher U_g (for example 6.4 cm/s).     

Gas Holdup and Volumetric Mass Transfer Coefficient in a Slurry Bubble Column

The gas holdup, ?, and volumetric mass transfer coefficient, k_La, were measured in a 0.051 m diameter glass column with ethanol as the liquid phase and cobalt catalyst as the solid phase in concentrations of 1.0 and 3.8 vol.‐%. The superficial gas velocity U was varied in the range from 0 to 0.11 m/s, spanning both the homogeneous and heterogeneous flow regimes. Experimental results show that increasing catalyst concentration decreases the gas holdup to a significant extent. The volumetric mass transfer coefficient, k_La, closely follows the trend in gas holdup. Above a superficial gas velocity of 0.04 m/s the value of k_La/? was found to be practically independent of slurry concentration and the gas velocity U; the value of this parameter is found to be about 0.45 s^–1. Our studies provide a simple method for the estimation of k_La in industrial‐size bubble column slurry reactors.     

Pressure and voidage fluctuations in slugging fluidized beds

In fully developed slugging fluidized beds, the maximum amplitude of absolute pressure fluctuations is reached with increasing superficial gas velocity when the slug length reaches a maximum, and the separation distance between successive slugs starts to decrease. U_c, the superficial gas velocity at which absolute pressure fluctuations reach a maximum, thus indicates the early stage of transition from slugging to core-annular flow. U_c, identified based on standard deviations of differential pressure fluctuations or local voidage fluctuations, can be predicted by slug flow models and does not signify a transition to turbulent fluidization.     

Transition velocity and bed expansion of two-phase (liquid-solid) fluidization systems

Hydrodynamic transition experiments for two-phase (liquid-solid), both upward and downward, liquid flow systems were performed in a 127-mm diameter column. The particles were 3.2-mm polymer (1,280 kg/m³), 5.8mm polyethylene (910, 930, 946 kg/m³), 5.5-mm polystyrene (1,021 kg/m³) and 6.0-mm glass (2,230 kg/m³) spheres, with water, aqueous glycerol solution and silicone oil as liquids. The dimensionless pressure gradient increases initially with increasing liquid velocity, but decreases gradually with increasing liquid velocity beyond U_lmf due to bed expansion. The non-dimensionalized pressure gradient using the liquid/solid mixture density increases with increasing liquid velocity and then reaches a constant value close to unity beyond U_lmf. The minimum fluidization Reynolds number for liquid-solid system increases with increasing Archimedes number including both heavier and lighter than the density of the liquid phase. U_lmf should be the same for both upward and downward fluidization systems since the Ergun equation is based on the main assumption that drag force of the superficial liquid velocity, U_lmf, is equal to the net difference between gravitational and buoyancy forces.     

Mass transfer and deodorization efficiency in a countercurrent spray tower for low superficial gas velocities

The aim of this study was to characterize mass transfer and deodorization efficiency in a countercurrent spray tower for low superficial gas velocities. The influence of operating parameters (U_G = 0.005 to 0.025 m s^?1, U_L = 6.1 × 10^?5 to 2.4 × 10^?4 m s^?1) on the liquid retention (ε_L), the drop diameter (d_g), the interfacial area (a) and the overall liquid and gas phase mass transfer coefficients (K_La, K_Ga) were estimated. The spray efficiency of some malodorous compounds was also estimated. A negative influence of the superficial gas velocity was demonstrated, during the spraying of water or chemical neutralizing scrubbing solutions. There was also an increase with the liquid flow rate. Abatements obtained were very good with respect to ammonia (>90%), and acceptable for the other compounds.     

On the onset of incipient fluidization

The onset of incipient fluidization is investigated theoretically and simulated by a computational fluid dynamics (CFD) procedure. The onset of incipient instability in a particle bed is preceded by stable gas diffusion in the interstices and is caused by a critical momentum force that may overcome the inertia of the particles. The critical momentum force is provided by the critical superficial gas velocity U_c in the form of critical mass flux of diffusion. It is found that the first movement of particles may be predicted by a critical transient Rayleigh number determined by a critical superficial velocity equals to the minimum fluidization velocity, U_mf. The onset of incipient fluidization was found to occur at a critical transient Rayleigh number of 3.1, which is close to the lowest theoretical value for buoyancy convection in a porous medium bounded by free surfaces. Consequently the onset times of incipient fluidization may be predicted accurately. The finding has been found to be supported by the present CFD study, past experiments and simulations in the literature.     

AGGLOMERATION BEHAVIOR IN A BUBBLING FLUIDIZED BED AT HIGH TEMPERATURE

Minimum fluidization velocity and agglomeration behavior were investigated at high temperature in an 80?×?30?mm two-dimensional quartz fluidized bed and in an 82?mm i.d. circular fluidized bed. Bed materials tested were two sizes of glass beads as well as three sizes of fluidized bed combustor (FBC) ash. The minimum fluidization velocity decreased with increasing bed temperature, whereas the minimum sintering fluidization velocity increased with the bed temperature. The sintering of glass beads belongs to visco plastic sintering, the first type. FBC ash agglomerate has higher amounts of SiO₂, Al₂O₃, Na₂O, K₂O, and SiO₂ than in the original ash, indicating that low melting eutectics were formed and that the liquid phase in a silicate system was formed. The agglomeration of FBC ash belongs to the second type, an excessive quantity of liquid being formed by melting or chemical reaction.     

Gas‐Liquid Mass Transfer in Fibrous Bed Reactor with Counter‐Current Liquid Recycle

In this work, the gas‐liquid mass transfer in a lab‐scale fibrous bed reactor with liquid recycle was studied. The volumetric gas‐liquid mass transfer coefficient, k_La, is determined over a range of the superficial liquid velocity (0.0042–0.0126 m.s^–1), gas velocity (0.006–0.021 m.s^–1), surface tension (35–72 mN/m), and viscosity (1–6 mPa.s). Increasing fluid velocities and viscosity, and decreasing interfacial tension, the volumetric oxygen transfer coefficient increased. In contrast to the case of co‐current flow, the effect of gas superficial velocity was found to be more significant than the liquid superficial velocity. This behavior is explained by variation of the coalescing gas fraction and the reduction in bubble size. A correlation for k_La is proposed. The predicted values deviate within ± 15 % from the experimental values, thus, implying that the equation can be used to predict gas‐liquid mass transfer rates in fibrous bed recycle bioreactors.     

Hydrodynamics of three-phase fluidized bed systems examined by statistical, fractal, chaos and wavelet analysis methods

Hydrodynamic studies were conducted in gas-liquid-solid systems (0.1 m ID, 2 m high) of 3.0 mm glass beads and of 2.1 mm polypropylene low-density particles, with particles densities of 2471 and 1290 kg/m³, respectively. Simultaneous measurement of differential pressure and bubble conductivity probe signals sampled at 500 Hz for 60 s enabled the investigation of the change in flow structure in relation to the flow regime transitions. Superficial gas velocities ranged between 0.010 and 0.052 m/s for polypropylene particles, and extended to 0.12 m/s for glass beads, while the superficial liquid velocities covered the ranges of 0.0007-0.045 m/s for polypropylene particles, and ranged up to 0.056 m/s for glass beads.Spectral analysis of the pressure fluctuations revealed a transition from dispersed to coalesced bubbling flow with decreasing liquid velocity for a given superficial gas velocity. The use of a conductivity probe facilitated characterization of the local flow structure in terms of bubble movement. The measurements were extensively analyzed using fractals and chaos, power spectra frequency analysis and wavelet decomposition in addition to the standard statistical analyses. The coefficient of variation of the bubble probe signals was found to be the most effective in deducing the transition velocity between coalesced and dispersed bubbling flow regimes, while wavelet energy confirmed the similarity in the distribution between two axial positions once operated in the dispersed flow regime. Comparison of the flow structure between glass beads and polypropylene particles showed that both the minimum liquid fluidization velocity and the transition velocity between the bubble flow regimes were much higher for the glass beads than for the lighter polypropylene particles. Furthermore, the standard deviations of the decomposed bubble probe signals through wavelet transformation successfully highlighted the difference between the two systems of particles.     

A consideration of circulating bed electrodes for the recovery of metal from dilute solutions

A comparison is made of three types of circulating particulate electrodes: spouted (circulating) bed (SBE), vortex bed (VBE) and moving bed (MBE). In applications such as metal recovery, all electrodes perform similarly in terms of current efficiency. On the basis of scale-up, it appears that the spouted bed electrode is the preferred system.Nomenclature I cell current (A) - F Faraday constant (94487 C mol^–1) - C dimensionless concentration - C _F friction factor - C ₀ Initial concentration (mol m^–3) - D pipe equivalent diameter (m) - e _b bed voidage - e _c voidage of conveying section - L bed length (m) - S _b cross section area of bed (m²) - S _T cross section area of conveying section (m²) - T dimensionless time=It/nFVC ₀ - U _f superficial liquid velocity in conveying (m s^–1) - U _i particle terminal velocity corrected for wall effects (m s^–1) - U _s particle velocity in transport (m s^–1) - U _SL slip velocity (m s^–1) - t time (s) - V electrolyte volume (m³) - V _f liquid velocity in the bed (m s^–1) - V _mf minimum fluidization velocity (m s^–1) - V _s particle velocity in the bed (m s^–1) - P pressure drop (NM^–2) - fluid density (kg m^–3) - _s particle density (kg m^–3) - Re Reynolds number     

Magnetic resonance imaging of gas–solid fluidization with liquid bridging

Magnetic resonance imaging is used to generate snapshots of particle concentration and velocity fields in gas–solid fluidized beds into which small amounts of liquid are injected. Three regimes of bed behavior (stationary, channeling, and bubbling) are mapped based on superficial velocity and liquid loading. Images are analyzed to determine quantitatively the number of bubbles, the bubble diameter, bed height, and the distribution of particle speeds under different wetting conditions. The cohesion and dissipation provided by liquid bridges cause an increase in the minimum fluidization velocity and a decrease in the number of bubbles and fast particles in the bed. Changes in liquid loading alter hydrodynamics to a greater extent than changes in surface tension or viscosity. Keeping U/U_mf at a constant value of 1.5 produced fairly similar hydrodynamics across different wetting conditions. The detailed results presented provide an important dataset for assessment of the validity of assumptions in computational models. © 2017 American Institute of Chemical Engineers AIChE J, 64: 2958–2971, 2018     

Gas‐Liquid Mass Transfer in a Slurry Bubble Column at High Slurry Concentrations and High Gas Velocities

The volumetric mass transfer coefficient k_La in a 0.1 m‐diameter bubble column was studied for an air‐slurry system. A C₉‐C₁₁ n‐paraffin oil was employed as the liquid phase with fine alumina catalyst carrier particles used as the solid phase. The n‐paraffin oil had properties similar to those of the liquid phase in a commercial Fischer‐Tropsch reactor under reaction conditions. The superficial gas velocity U_G was varied in the range of 0.01 to 0.8 m/s, spanning both the homogeneous and heterogeneous flow regimes. The slurry concentration ?_S ranged from 0 to 0.5. The experimental results obtained show that the gas hold‐up ?_G decreases with an increase in slurry concentration, with this decrease being most significant when ?_S < 0.2. k_La/?_G was found to be practically independent of the superficial gas velocity when U_G > 0.1 m/s is taking on values predominantly between 0.4 and 0.6 s^–1 when ?_S = 0.1 to 0.4, and 0.29 s^–1, when ?_S = 0.5. This study provides a practical means for estimating the volumetric mass transfer coefficient k_La in an industrial‐size bubble column slurry reactor, with a particular focus on the Fischer‐Tropsch process as well as high gas velocities and high slurry concentrations.     

Determination of minimum fluidization velocity by pressure fluctuation measurement

Using the standard deviation of pressure fluctuations to find the minimum fluidization velocity, U_mf, avoids the need to de-fluidize the bed so U_mf, can be found for operational bubbling fluidized beds without disrupting the process provided only that the superficial velocity may be altered and that the bed remains in the bubbling fluidized state. This investigation has concentrated on two distinct aspects of the pressure fluctuation method for U_mf determination: (1) the minimum number of pressure measurements required to obtain reliable estimates of standard deviation has been identified as about 10000 and (2) pressure fluctuation measurements in the plenum below the gas distributor are suitable for U_mf determination so the problems of pressure probe clogging and erosion by bed particles may be avoided.