Fluidization Characteristics of SiO2 Nanoparticles in an Acoustic Fluidized

Under the action of an acoustic field, the fluidization behavior of 5–10 nm SiO₂ nanoparticles, with and without surface modification, was investigated. In a packed bed, the sound wave energy has a significant influence on the compact ratio of the bed. Experimental results indicated that the bed of nanoparticle agglomerates can be fluidized smoothly with the assistance of an acoustic field, and the minimum fluidization velocity is initially reduced dramatically with increasing sound frequency and then rises with increasing sound frequency. Under the same experimental conditions, the minimum fluidization velocity of 5–10 nm SiO₂ nanoparticles is greater than that of 5–10 nm SiO₂ nanoparticles with surface modification. The collapse of the bed demonstrates that SiO₂ nanoparticles, surface modified using organic compound, have longer minimum collapse times than SiO₂ nanoparticles.     

Hydrodynamic characteristics of cold-bed circulating fluidized beds for the methanol to olefins process

The effects of the riser inlet velocity (2.2–3.9 m/s), seal-pot inlet velocity (2.4–7.1 U _mf ), aeration flow rate (2.5×10^?7–3.7×10^?6 m³/s) in seal-pot, and solid inventory (0.15–0.2 kg) on the hydrodynamic characteristics of a 9 mm-ID×1.9 m-high cold-bed circulating fluidized bed for methanol to olefins (MTO) process were investigated. FCC (Engelhard; 82.4 μm) particles were used as bed particles. Most of the experimental flow regimes were observed in fast fluidization and pneumatic transport regimes. The axial solid holdup in a riser increased with increasing solid mass flux and solid inventory. Solid mass flux increased proportionally until reaching a maximum value and then decreased with increasing seal-pot inlet velocity. The obtained hydrodynamic characteristics in the cold-bed circulating fluidized beds were compared with previous results.     

Liquid hold-up and minimum fluidization velocity in a turbulent contactor

Transient-response experiments have been performed in conjunction with bed expansion measurements to determine liquid hold-ups and minimum fluidization velocities in a 12-in. turbulent-bed gas-liquid contactor. The amount of liquid hold-up was found to be independent of gas velocity, but dependent upon both liquid rate and packing diameter in the same manner as reported for conventional fixed-bed absorbers. The data on minimum fluidization velocity, G_mf, which was interpreted in the present study as the maximum gas mass velocity at which the bed maintained its static height, showed a considerable variation with packing diameter, d_p and liquid flow rate, L. A correlation of G_mf with d_p and L was presented.     

CFD simulation of flow pattern and jet penetration depth in gas-fluidized beds with single and double jets

A simple two-fluid model is validated by comparing single-jet fluidization experiments and numerical predictions. Subsequently, flow pattern and jet penetration depth are explored numerically in the bed with double jets under equal and unequal gas velocities. Glass balltoni with a density of 2550 kg/m³ and a diameter of 275 μm is employed as solid phase. The model used in this study considers the effect of the dispersed solid phase on both gas and particle momentum equations of the inviscid model A (Gidaspow, 1994). Numerical simulations are carried out in the platform of CFX 4.4, a commercial CFD code, together with user-defined FORTRAN subroutines. Both jet penetration depth and jet frequency predicted are in good quantitative agreement with measurements in an incipiently fluidized bed with a single jet. By combining solid volume fraction distribution and particle-phase velocity vector profile, three flow patterns (isolated, merged and transitional jets) are identified in the gas-fluidized bed with double jets, which depend more on the nozzle distance than the jet gas velocity. For the equal jet gas velocity, the jet penetration depth decreases with increasing nozzle distance in the merged-jet and transitional-jet regions, then reaches a minimum value in the transitional-jet region, and finally keeps steady in the isolated-jet region. For the unequal jet gas velocity, the merged jet penetration depth increases with increase in the velocity of one jet as the other jet gas velocity is fixed, whilst the jet penetration depths change a little in the transitional-jet region and remain a constant in the isolated-jet region.     

Effects of pressure and type of gas on particle-particle interaction and the consequences for gas—solid fluidization behaviour

The fluidization behaviour of cracking catalyst has been studied up to pressures of 15 bar with different fluidization gases (Ar, N₂, H₂). A number of parameters of both the homogeneous and heterogeneous fluidized bed has been examined experimentally.The experimental results reveal that the minimum fluidization velocity (U_mf) is independent of the pressure. The bubble point velocity (U_bp) and the maximum bed expansion (H_bp) at this velocity increase with increasing pressure. This also holds for the dense phase voidage (ε_d) and the dense phase gas velocity (U_d) in the bubbling bed. The bubble size decreases drastically with increasing pressure. However, the above-mentioned parameters are also strongly dependent on the type of fluidization gas used.The cohesion constant of the powder was measured, using a tilting bed technique. The results reveal that the cohesion constant increases with increasing pressure. Analysis of the results of adsorption measurements of the different gases to the solid reveals for the adsorption as well as for the cohesion and for the beu expansion the same pressure dependence.It is believed that the gas adsorption influences the cohesion between the particles and hence the elasticity modulus introduced by Rietema and Mutsers [1,2]. The increasing elasticity modulus with increasing pressure also explains the increasing bed expansion with pressure.     

Fluidization enhancement of agglomerates of metal oxide nanopowders by microjets

The quality of gas–solid fluidization of agglomerates of nanoparticles has been greatly enhanced by adding a secondary flow in the form of a high‐velocity jet produced by one or more micronozzles pointing vertically downward toward the distributor. The micronozzles produced a jet with sufficient velocity (hundreds of meters per second), turbulence, and shear to break‐up large nanoagglomerates, prevent channeling, curtail bubbling, and promote liquid‐like fluidization. For example, Aerosil R974, an agglomerate particulate fluidization (APF) type nanopowder, expanded up to 50 times its original bed height, and difficult to fluidize agglomerate bubbling fluidization (ABF) type nanopowders, such as Aeroxide TiO2 P25, were converted to APF type behavior, showing large bed expansions and homogeneous fluidization without bubbles. Microjet‐assisted nanofluidization was also found to improve solids motion and prevent powder packing in an internal, is easily scaled‐up, and can mix and blend different species of nanoparticles on the nanoscale. © 2009 American Institute of Chemical Engineers AIChE J, 2010     

喷嘴进料对提升管进料段内颗粒浓度分布的影响

在提升管冷模实验装置上考察了喷嘴进料对颗粒浓度径向分布的影响规律. 结果表明,提升管进料段内存在3种形式的颗粒浓度径向分布,在距喷嘴较近的轴向区域,颗粒浓度沿径向呈明显的W形分布,喷嘴进料对颗粒流动的影响很强;在距喷嘴较远的轴向位置,颗粒浓度沿径向呈环-核分布,喷嘴进料对颗粒流动的影响很弱;在二者之间,颗粒浓度沿径向呈弱W分布,喷嘴进料对颗粒流动具有一定影响. 随着喷嘴气速的增加或预提升气速的减小,颗粒浓度逐渐由W形分布转变为环-核分布,喷嘴进料对颗粒流动的影响逐渐减弱. 采用喷嘴射流动量与预提升来流动量比Mj/Mr考察了操作参数及装置结构尺寸对提升管进料段内颗粒浓度径向分布的综合影响. 在实验范围内,动量比对进料段内颗粒浓度径向分布及颗粒流动行为具有明显的影响规律,随着动量比的增加,颗粒浓度逐渐由W形分布转变为环－核分布,操作参数及装置结构尺寸对颗粒流动的影响逐渐减小. 在动量比小于4.21时,操作参数及装置结构尺寸对颗粒流动的影响在H=0.675~1.075 m间的轴向位置基本结束;在动量比增大为4.21时,操作参数及装置结构尺寸对颗粒流动的影响在H=0.375~0.675 m间的轴向位置便已基本结束.     

Tomographic diagnosis of gas maldistribution in gas-solid fluidized beds

Phase distribution is one of the key hydrodynamic parameters useful for the design and performance assessment of fluidized bed dryers (FBDs). It has direct influence on the drying rate, thermal efficiency, residence time distribution and degree of mixing. The quality of fluidization strongly depends on the uniformity of distribution of the fluidizing gas and the physical properties of the material to be fluidized. In the present work, gamma ray tomography (GRT) study was carried out in the form of chordal solid hold-up, which was found to be greatly influenced by the gas distributor design. The performance of a gas distributor due to the prevalent practice of operating at lower values of distributor-to-bed pressure drop ratio was characterized in a 0.15 m diameter fluidized bed dryer over a broad range of superficial gas velocity. The effects of various parameters such as solids loading, particle size and particle density were analyzed with the help of the reconstructed solid hold-up profiles. The fluidization was studied in terms of maldistribution factor (χ), a value of 5% or less can be obtained by properly designing distributor for a given bed loading, particularly for batch fluidized bed dryers. An industrial size fluidized bed dryer of 1 m diameter was also examined tomographically to obtain quantitative information on the solid hold-up distribution within the bed.     

Local solid mixing in gas–solid fluidized beds

Diffusivity of the solid particles in a 152-mm ID gas–solid fluidized bed was determined at different regimes of fluidization. The gas was air at room temperature and atmospheric pressure and the solids were 385 μm sand or 70 μm FCC particles. The experiments were done at superficial gas velocities from 0.5 to 2.8 m/s for sand and 0.44 to 0.9 m/s for FCC (in both bubbling and turbulent regimes). Movement of a tracer was monitored by radioactive particle tracking (RPT) technique. Once the time-position data became available, local axial and radial diffusivity of solids were calculated from these data. Calculated diffusivities are in the range of 3.3×10⁻³ to 5.6×10⁻² m²/s for axial and 2.6×10⁻⁴ to 1.5×10⁻³ m²/s for radial direction. The results show that the diffusivities, both axial and radial, increase with superficial gas velocity and are linearly correlated to the axial solid velocity gradient. Solid diffusivity in a bed of FCC was found to be higher than that of a bed of sand at the same excess superficial gas velocity.     

A comparative study of a fluidised bed reactor and a gas lift loop reactor for the IBE process. Part II: Hydrodynamics and reactor modelling

A fluidised bed reactor with liquid recycle (FBR) and an external loop gas lift reactor (GLR) were designed for the production of isopropanol—butanol mixtures by immobilised Clostridium spp. and scaled down to laboratory scale (part I). Hydrodynamic models were set up for the two laboratory scale reactors. Liquid mixing in the 10 dm³ FBR was described by 10 tanks in series. Fluidisation velocities, bed expansions and axial dispersion coefficients agreed well with literature data. Liquid mixing in the 15 dm³ GLR was described by 100 tanks in series. The gas hold-up and circulation velocity were found to decrease with increasing hold-up of solids, in accordance with literature indications. No influence of the hold-up of solids on the axial dispersion coefficient was determined. An integrated reactor model was set up for both reactors, using the hydrodynamic and kinetic model. Actual fermentation data are presented and compared with model predictions in part III of this study; this part will also include a comparison of reactor performances and scale up aspects.     

气固流化床内射流穿透深度的CFD模拟及其实验验证

在经典的Gidaspow无黏性双流体模型中考虑离散颗粒对流体和固体动量守恒方程的影响后,建立了一个具有模拟大规模流化床内气固两相流体动力学特性潜在优势的简化数学模型。在CFX4.4商业化软件平台上通过增加用户自定义子程序考察了二维气固流化床(高2.00 m、宽0.30 m)内射流气速、喷嘴尺寸、环隙气速和静床高度对射流穿透深度的影响,并以树脂颗粒(粒径670 μm、密度1474 kg·m^-3)为研究对象在厚度为0.025 m的矩形床内进行了对比实验。结果表明,选取空隙率为0.8的等高线作为射流边界比较合适;射流穿透深度随射流气速或射流喷口尺寸的增加而增大;射流周围环隙气速由0变到最小流化速度时,射流穿透深度随环隙气速增加而增大,在最小流化速度时达到最大值,然后随环隙气速增加单调减小,当环隙气速大于2.5倍最小流化速度时,射流穿透深度减小程度变缓;在相同射流气速下射流穿透深度随着静床高度的增加而减小,静床高度对射流穿透深度的影响随着射流气速增加呈现扩大的趋势。     

Characteristics of solid hold up and circulation rate in the CFB reactor with 3-loops

The effects of the U_o, PA/[PA+SA] ratio, total solid inventory and fluidizing velocity of loopseal on the axial solid holdup and the solid circulation rate have been determined with different particle sizes (174, 199, 281, 377 μm) and particle types (silica sand: narrow PSD, coal ash: wide PSD) in a CFB reactor with 3-loops. A simple model for solid hold-up based on the previous works was in agreement with the experimental data. With increasing U_o, G_s increased exponentially, and in the center-loop, G_s was 1.5 times larger than that found in the other side-loops. As the PA/[PA+SA] ratio increased, and as SA injection port was placed at a lower part in the riser, the axial solid holdup and G_s increased. With increasing fluidizing velocity of loopseal to about 1.5u_mf, G_s somewhat increased, but above the gas velocity of 1.5u_mf, the loopseal lost the ability of the control ofGs. The following correlation for the solid circulation rate in the CFB was developed with good accuracy; G, = ϕ,[PA/TA]²⁺[H _l /H _l ]^0.5[Ar]^-188[Fr]^2k+[KU₁/U _l ]^3-45     

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.     

Numerical Simulation of Liquid–Solid Countercurrent Fluidization inside an Extraction Column Based on Particle Trajectory Model

The liquid–solid countercurrent fluidization process in an extraction column was numerically simulated based on the particle trajectory model of Eulerian–Lagrangian method. The simulation approach was validated by previous experiments. A power function correlation was proposed for dimensionless slip velocity U_slip/U_t and hold-up fraction ϕ, and the operational zone in the countercurrent fluidization was determined. Simultaneous countercurrent fluidization of particles with different diameters was also simulated. The comparison shows that the simulation results are consistent with the calculation values from the multi-particle free sedimentation model based on non-interference assumption, verifying the reliability of the approach in present work.     

Bubbling and Bed Expansion Behavior Under Vibration in a Gas‐Solid Fluidized Bed

The effect of vibration on the flow patterns and fluidization characteristics including the minimum fluidization velocity (u_mf), the void fraction (ϵ_mf) at u_mf and the bed expansion ratio were examined. The powders used were spherical glass beads and their diameters were 6, 20, 30, 60 and 100μm. For group A powders, the manner in which the vibration affects the bubble formation was examined from the bed expansion ratio and the index of n/4.65. The area of the homogeneous fluidization region was also observed. The homogeneous fluidization region was broadened at a certain vibration strength, where the value of n/4.65 was a minimum. The bubble formation was observed even for 20μm powder (group C), at large vibration strengths and at high gas velocities. Under such conditions, the bed expansion ratio increased suddenly due to bubble formation. The bubbles broke the irregular bed structure, including various properties of agglomerates. Although the channel breakage was dominant flow pattern for group C powders, the bubbles also played an important role in the improvement of the 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.     

Rotating fluidized bed with a static geometry: Guidelines for design and operating conditions

Computational fluid dynamics (CFD) simulations of the hydrodynamic behavior of rotating fluidized beds in static geometry (RFB-SG) are carried out for gas–solid flows. The rotating motion of the reactor bed is induced by the tangential injection of the gas along the circumference of the fluidization chamber. Steep gradients in the gas velocity fields both in radial and tangential direction generate turbulence. The radial and tangential drag forces fluidize the particle bed in both radial and tangential direction.An Eulerian two-fluid model is used. Gas phase turbulence is accounted for by a k–ε model adapted for rotational flows. The RFB-SG simulations provide guidelines for a design and operation with a high efficiency in gas–solid momentum transfer, excellent gas–solid separation and limited solids losses. Hydrodynamic variables like the centrifugal force, the injection pressure, the radial and tangential slip velocities, solids hold-up are calculated for both polymer particles (300 μm, 950 kg/m³, Geldart Group B) and glass beads (70 μm, 2500 kg/m³, Geldart Group A) to allow for a comparison among different fluidization chamber designs. Unstable bed behavior, like slugging and channeling, is also numerically predicted.     

Heat transfer from a grid jet in a large fluidized bed

Heat transfer from a vertical grid jet within a 2 ft diameter and 4 ft deep fluidized bed of cracking catalyst was studied. The test nozzle diameter was varied from ¼ to 1 in. and the nozzle velocity from 50 to 250 ft /sec which is within the range of industrial practice. The axial temperature data have been related to a Froude, a Reynolds and a Nozzle number: In (δT/δT_o) = –58.1 Fr^?0.562 No^1.08 Re^?0.112 A simple jet quenching model yielded heat transfer coefficients between the fluid bed and grid jet which ranged from 300 to 1200 Btu/ft² hr.^o F.     

Application of discrete element method simulation for studying fluidization of nanoparticle agglomerates

The usefulness of discrete element method simulation for studying fluidization of nanoparticle agglomerates is explored. Nanoparticle agglomerates were simulated by using solid particles of equivalent sizes and densities. Validity of the present simulation was assessed through comparisons of simulation results and experimental observations of bed expansion, characteristic fluidization behaviour, and dense‐bed settling. The simulation was then used to investigate initial bed expansion and bed uniformity under particulate fluidization conditions. The role of inter‐agglomerate interparticle force in fluidization of nanoparticle agglomerates was examined. A stability analysis originally developed for addressing the transition from particulate to bubbling fluidization for conventional particles was used for predicting the start of bubbling in fluidized beds of nanoparticle agglomerates.