PRODUCTION OF SUB-MICROMETER SIZE O/W EMULSIONS USING A HIGH-PRESSURE WET-TYPE JET MILL

Emulsions composed of sub-micrometer size droplets were obtained by a high-pressure wet-type jet mill. Liquid paraffin as dispersed phase, aqueous sucrose solution as continuous phase, and anionic sodium dodecyl sulfate as emulsifying agent were used as the model emulsification system. Droplet size distribution, Sauter mean diameter (d₃₂), geometric standard deviation of the droplet size distribution (σ_g), and emulsion viscosity (η_e) were investigated under the various combinations of process variables: dispersed phase volume fraction (Φ), dispersed phase viscosity (η_d), continuous phase viscosity (η_c), processing pressure, and number of passages through the wet-type jet mil. d₃₂ and σ_g were correlated with the process variables. For the entire experimental range, maximum droplet diameter was varied with d₃₂ in the range of 1.7-1.9-fold, and a correlation was proposed with K (=η_d/η_c) as a variable. Above a marginal value of d₃₂, at Φ ≥ 0.1, emulsions exhibited Newtonian behavior and could be described well by the Yaron and Gal-Or model. Below the marginal values, emulsions strongly depended on d₃₂ and exhibited a shear-thinning behavior. The relation between η_e and d₃₂ for such emulsions was modeled by use of particle Reynolds number.     

EVALUATION OF CONDUCTIVE HEAT TRANSFER MECHANISMS BETWEEN AN IMMERSED SURFACE AND THE ADJACENT LAYER OF PARTICLES IN BUBBLING FLUIDIZED BEDS

The evaluation of the heat transfer coefficient h_wp between a heat exchanging surface immersed in a gas fluidized bed and the adjacent layer of dense phase particles is analyzed in this contribution. Gas convective and radiant effects are not included in the present analysis.

The inclusion of h_wp, or an equivalent formation, in mechanistic models describing heat transfer has been necessary because the sudden voidage variation close to the immersed wall restrains significantly the heat transfer rate. However, there is not at present a widely accepted expression to evaluate h_wp.

A precise formulation for h_wp accounting for transient conduction inside spherical particles, the Smoluchowski effect, the concentration of particles in the adjacent layer (N_p) and an effective separation gap (l₀) is developed here.

Although N_p can be estimated, in principle, from experimental evidence in packed beds, and it is reasonably expected that l₀ = 0, the analysis of experimental heat transfer rates in moving beds, packed beds, and bubbling fluidized beds indicate that values of h_wp are, in general, smaller than expected from these assumptions. Appropriate values of l₀ and N_p are then stimated by fitting the experimental data.

The probable effect of surface asperities is also discussed by analyzing a simplified geometrical model. It is concluded that the parameter l₀ can be also effective to account for particle roughness, independently of thermal properties.     

RADIAL DISPERSION AND BUBBLE CHARACTERISTICS IN THREE-PHASE FLUIDIZED BEDS

The effects of gas and liquid velocities, liquid viscosity and particle size on the radial dispersion coefficient of liquid phase (D_r) and the bubble properties in three-phase fluidized beds have been determined. A new flow regime map based on the drift flux theory in three-phase fluidized beds has been proposed.

In three-phase fluidized beds, D, increases with increasing gas velocity in the bubble coalescing and in the slug flow regimes, but it decreases in the bubble disintegrating regime. The coefficient exhibits a maximum value in the bed of small particles with increasing liquid velocity at lower gas velocities. However, it increases with increasing liquid velocity at higher gas velocities. In two and three-phase fluidized beds of larger particles (6,8 mm), D_r exhibits a maximum value with an increase in liquid viscosity at lower gas velocities, but it increases at higher gas velocities. The mean bubble chord length and its rising velocity increase with increasing gas velocity and liquid viscosity. However, the bubble chord length decreases with an increase in liquid velocity and it exhibits a maximum value with increasing particle size in the bed. The radial dispersion coefficients in the bubble coalescing and disintegrating regimes of three-phase fluidized beds in terms of the Peclet number in the present and previous studies have been well represented by the correlations based on the concept of isotropic turbulence theory.     

HIGH TEMPERATURE INCIPIENT FLUIDIZATION IN MONO AND POLYDISPERSE SYSTEMS

Tests have been made on the behaviour of fluidized beds at high temperature (15-950^°C). Bed materials used were silica sand of different sizes.

Bed voidage at minimum fluidization conditions was found to be dependent on temperature rise and on increase in Reynolds number. Plots of bed voidage function, bed voidage and Ar versus Re_mf show a change in the behaviour at Reynolds numbers between approximately 0.75-2. This is explained in terms of a variation in the fluid flow pattern inside the bed: at very low Re_mf creeping flow exists, but at higher values of Reynolds number, separation of boundary layer takes place and a wake appears at the rear of the particle, creating a low pressure zone. This contributes to attractive forces between particles at the minimum fluidization conditions, decreasing the value of ε_mf. If Re_mf increases, the separation point moves towards the rear of the particle and the wake shrinks; attractive forces decrease, and ε_mf increases.

Data on u_mf, both for monodisperse systems and binary mixtures, are compared with predictions from different equations.     

MULTIPHASE HYDRODYNAMICS AND SOLID-LIQUID MASS TRANSPORT IN AN EXTERNAL-LOOP AIRLIFT REACTOR—A COMPARATIVE STUDY

The solid-solid mass transfer performance of an external-loop airlift reactor was measured by dissolution of benzoic acid coated on nylon-6 particles, and the hydrodynamics of the gas-liquid-solid multiphase system in the airlift reactor were investigated. The solid-liquid system was designed to simulate the micro-carrier culture of animal cells, and some typical suspensions of immobilized enzyme particles.

The solid-liquid mass transfer coefficient remained constant below a superficial air velocity of 0.04 ms^-1 for the particles examined, but increased rapidly with further increase in gas velocity. Solids loading (0.3-3.5% w/w) did not affect the mass transfer coefficient in turbulent flow.

The mass transfer coefficient was correlated with energy dissipation rate in the airlift reactor. The mass transfer coefficient in stirred vessels, bubble columns, fluidized beds, and airlift reactors was compared.

Over an energy dissipation Reynolds number of 4-400, the solid-liquid mass transfer coefficient in the airlift device was comparable to that obtainable in fluidized beds. The performance of the airlift was distinctly superior to that of bubble columns and stirred tanks.     

GAS MIXING IN SLUGGING AND TURBULENT FLUIDIZED BEDS

The gas phase mixing in a fluidized bed of glass beads (d_p = 0.362 mm) in the slugging and turbulent flow regimes has been studied in a 0.1 m-ID × 3.0 m high Plexiglas column.

The gas dispersion in the downstream of the bed has been described by a diffusion process with the axial and radial dispersion coefficients. The radial dispersion coefficient of the gas phase is nearly constant with the variation of gas velocity in the slugging flow regime, but it increases with an increase in gas velocity in the turbulent flow regime.

Appreciable backmixing of the gas phase is pronounced in the slugging flow regime whereas the lower gas backmixing is produced in the turbulent flow regime. The gas backmixing coefficient increases with an increase in gas velocity in the slugging flow regime, but it decreases slightly with an increase in gas velocity in the turbulent flow regime.

The radial mixing and backmixing coefficients of the gas in terms of Peclet numbers have been correlated with the relevant dimensionless parameters (U_g/U_mf, p_s/p_g, d_p/D_t).

The gas flow pattern in the bed has been well represented by a simplified model based on the two gas phases in the dilute and dense phases which are percolating through the bed in plug flow. The present model can predict the gas exchange coefficient between the phases, the fractions of the dilute phase, the interstitial gas in the dense phase, and the interstitial gas velocity in the bed.     

Hydrodynamic characteristics of a circulating fluidized bed

In a circulating fluidized bed (7.8 cm-ID x 260 cm-high), flow regime of coal-air system at room temperature has been determined. Bituminous coal particles used were either 0.73 mm or 1.03 mm in the mean diameter having density of 1400 Kg/m³. The transition velocities from bubbling to turbulent beds and the transport velocities between turbulent and fast beds have been determined. The resulting transition velocities between bubbling and turbulent beds were 103 cm/s for 0.73 mm and 130 cm/s for 1.03 mm coal particles, respectively. The transport velocities between turbulent and fast beds were 180 and 209 cm/s for 0.73 and 1.03 mm particles, respectively. In addition, chocking velocities were determined at different solid feeding rates. The resulting values were in the range of 2.55-2.65 m/s for 0.73 mm particle and of 2.77-2.84 m/s for 1.03 mm particle, respectively. The published literature data of the transition velocity between bubbling and turbulent bed have been correlated with particle properties.     

FIBRE OPTIC AND CAPACITANCE PROBES IN TURBULENT FLUIDIZED BEDS

In the present study, local flow properties are investigated in pilot plant scale fluidized bed reactors using both fibre optic and capacitances probes. Measurements are conducted at ambient as well as at 150°C. The system used is air and spent FCC particles (mean particle diameter: 65pm). The static bed height is 1.6 m. Bubbling and turbulent regimes (V = 0.40 and 0.70 m/s) are investigated in two 0.3 and 0.5 m ID columns.

Bubble fraction under the bubbling regime and at room temperature, measured using fibre optic and capacitance probes, are in good agreement. However, in the turbulent regime, fibre optic probes are prone to underestimate the bubble fraction while capacitance sensors tend to overestimate it. These discrepancies between fibre optic and capacitance measurements increase with temperature.

Using capacitance probes, a prevalent flat bubble rise velocity profile is measured. This is assigned to the relatively slow response, to the size and to the geometry of the capacitance probes. Overall this gives an underestimation of the bubble frequency and an overestimation of the bubble contact time and the bubble contact length.

Due to the high fibre optic probe sensitivity, care should be taken in the interpretation of signals. Overestimation of bubble frequency leads to underestimation of both bubble contact times and bubble contact lengths.     

BUBBLE COMPUTATION, GRANULAR TEMPERATURES, AND REYNOLDS STRESSES

Bubbles were simulated in a two-dimensional fluidized bed with a constant inlet velocity using two computer codes, the IIT code and the MFIX code. The computational results were compared to the Jung et al. (2005) experiments in a thin bubbling bed of 530 μm glass beads. The use of higher order numerics produces better bubble resolution due to smaller numerical diffusion. The computed bubble sizes and their distributions agreed with the experiments. The simulations show that there is no bubble formation for sufficiently elastic particles.

CFD computations and previous experiments show that in the bubbling fluidized beds there exist two random oscillations. The first kind is due to random oscillations of particles and is measured by the conventional granular temperature. The second one is due to motion of bubbles and gives rise to Reynolds type stresses. It is shown that the particle granular temperature is much smaller than the bubble-like granular temperature computed from the average of the normal Reynolds stresses, measured by Cody using a shot noise technique.     

EXPERIMENTAL OBSERVATIONS OF FLUIDIZED BEDS AT ELEVATED PRESSURES

The object of the work described here was to elucidate the effects of operation under pressure on the physical behaviour of gas fluidized beds. Extensive measurements of various bubble properties such as size, shape and rise velocity in beds of coarse powders (mean particle diameters of 184 μm and 450μm) operated at pressures of up to 81 bar were made by photographing the images created by irradiation of the bed with X-rays, and analysing the bubble silhouettes thereby obtained. Most of the results presented here are averages of some 200 individual measurements.

Experimental evidence to support the following picture of the effect of pressurization on the behaviour of freely bubbling gas fluidized beds is presented. Both bubble interaction (tendency to coalesce) and the incidence of bubble splitting increase with increasing pressure; the two are intimately connected. The nett results are a decrease in bubble size with increasing pressure over most of the pressure range and an increase in the tendency for bubbles to distribute non-uniformly in a radial direction. This latter tendency probably causes gross solids circulation in the bed, and this in turn leads to higher bubble rise velocities than those observed for single bubbles under similar conditions. The splitting mechanism accounting for the decrease in bubble size was found to be intrusion of the wake into the bubble void by the flow of gas through the wake region of a leading bubble during pair coalescence.

An updated review of other published work relating to the subject of experimental observations of the effects of pressure on gas fluidized beds is included in the form of a table.     

RADIAL FLOW HOLLOW FIBERREVERSE OSMOSIS: EXPERIMENTS AND THEORY

The performance of a hollow fiber reverse osmosis system is studied both theoretically and experimentally. Experiments were carried out for applied pressure ranging from 200 to 400 psig, feed rates varying from 75 to 380 cc/sec and for feed concentrations up to 34,000 ppm of sodium chloride.

A mathematical model is proposed to predict productivity, ϕ, and product concentration, θ_p. The model involves solving membrane transport equations simultaneously with the hydrodynamic equations. The solubility-diffusion-imperfection, or pore diffusion model, is used to describe solute and solvent transport across the membrane. The axial gradients of shell side concentration, neglected in previous investigations, are taken into account. The differential equations are solved numerically by the 4th Order Runge-Kutta method.

Predicted values of ϕ and θ_p agree within 8% and 17% respectively, with experimental data over the entire range of operating conditions. However, membrane transport coefficients were found to be concentration dependent.

An approximate analysis shows that the concentration polarization is negligible in present day hollow fiber systems.     

DISPERSION IN TURBULENT FILM FLOW

Dispersion of miscible solutes from non-uniformly distributed instantaneous sources in two dimensional smooth open channel turbulent flows is described satisfactorily by using generalized dispersion theory wherein the coefficients of the dispersion model depend on time. The truncation of the generalized dispersion model after the first two terms is justified by comparison of results with experimental data. The behavior of the convective coefficient, k₁, (t) and the dispersion coefficient, k₂(t), for turbulent flow is studied in detail for all values of time, k₁(t) and k₂(t) reach asymptotic constant values at t≈ 0.5 y²_m,/ε at which time the generalized dispersion model reduces to the Taylor constant coefficient model.

Area mean concentration distributions predicted by the generalized dispersion theory agree with the experimental data of Fischer (1966) rather closely and the theoretical results agree with the data much better than those from Taylor's model.     

Drying Suspensions in a Pulsed Fluidized Bed of Inert Particles

The drying of carbohydrate suspensions on polypropylene particles in a pulsed fluidized bed was studied by means of a 2⁵ experimental design, to determine the effect of the air flow and temperature, suspension flow rate, and free section and rotating speed of the rotary plate on the Nusselt number, the moisture content of the product, and the percentage of solids retained inside the bed (which were minimized to 4.9 and 14.4%, respectively) with an air flow of 600 m³/h at 90°C and 720 mmHg, a suspension flow rate of 6 L/h, and a plate with 6% free section, rotating at 50 rpm.

Additionally, the effects of temperature, air flow, and suspension flow rate on the residence time distribution (RTD) were determined, using the stimulus-response methodology. The RTD was represented by 1.1 to 2 tanks in series, according to this model. The mean residence time of the dried carbohydrate particles was between 5.4 and 8.2 min.

Finally, an egg suspension could be dried at 4 L/h, with air at 90°C, with a mean residence time about 50% longer that that found for drying carbohydrate suspensions.     

EVALUATION OF OVERALL HEAT TRANSFER RATES BETWEEN BUBBLING FLUIDIZED BEDS AND IMMERSED SURFACES

A mechanistic model to evaluate heat transfer rates between the dense phase of gas fluidized beds and immersed surfaces has been recently presented by the authors. This model, denoted Generalized Heterogeneous Model (GHM), is formulated in terms of effective thermal properties for particles and interstitial gas. It has been conceived with the purpose of achieving a generalized formulation accounting simultaneously for conductive, gas convective and radiant effects.

The model was previously tested as regards its capability to predict radiative heat transfer rates in beds at high temperature and gas convective contribution in beds of large particles and high operating pressures.

It is the principal object of this contribution to evaluate the performance of the GHM for a wide range of particle sizes, covering from The purely conductive regime to the gas convection dominant regime.

Also, the main assumptions incorporated in the model are revised and some modifications are introduced, mainly on the basis of the results obtained by Mazza et al. (1997b).     

EFFECT OF IMPELLER DESIGN ON LIQUID PHASE MIXING IN MECHANICALLY AGITATED REACTORS

Liquid phase mixing time (θ_mix) was measured in mechanically agitated contactors of internal diameter 0.57 m, 1.0 m and 1.5 m. Tap water was used as the liquid phase. The impeller speed was varied in the range of 0.4-9.0 r/s. Three types of impellers, namely disc turbine (DT), pitched blade downflow turbine (PTD) and pitched-blade upflow turbine (PTU) were employed. The ratio of impeller diameter to vessel diameter (D/T) and the ratio of impeller blade width to impeller diameter (W/D) were varied over a wide range. The effects of impeller clearance from the tank bottom (C), the blade angle (φ), the number of blades (n_b), the blade thickness (k) and the total liquid height (H/T) were studied in detail. Mixing time was measured using the conductivity method.

Mixing time was found to have a strong dependance on the flow pattern generated by the impeller. Mixing time was found to decrease by decreasing the impeller clearance in the case of DT and PTU. However in the case of PTD it increases with a decrease in the impeller clearance. Similar trend of the effect of impeller clearance on θ_mix, was observed for all the other PTD impellers with different diameter, number of blades and blade angle (except 60^° and 90^°). All the impeller designs were compared on the basis of power consumption and on this basis optimum design recommendations have been made. For PTD impellers, a correlation has been developed for the dimensionless mixing time.     

GAS HOLDUP IN UNIFORMLY AERATED BUBBLE COLUMN REACTORS

Conditions of the stable performance of gas distributing plates were studied and the effect of plate geometry on gas holdup of uniformly aerated gas-liquid beds was investigated. The ratio of plate holes opened for gas passage was determined as a function of gas hole velocity and critical values of gas hole velocity corresponding to the onset of stable performance of distributing plates were obtained.

Two regimes of bubbling were observed under conditions of stable uniform gas distribution; the regions of their existence being determined by the values of gas flow velocity and distributing plate parameters. Considerable increase of gas holdup was observed in the region of “foam” bubbling compared to the “turbulent” bubbling regime commonly encountered in bubble column reactors. The character of the bed and hence its gas holdup value were affected by the geometry of distributing plates in the “foam” bubbling region while no such effect was observed under “turbulent” bubbling conditions.     

Detailed measurements of flow structure inside a dense gas–solids fluidized bed

Detailed local flow structures are investigated in bubbling and turbulent fluidized bed with FCC particles. The operating conditions ranges from 0.06 to 1.4 m/s. Extensive experiments are carried out using a newly developed optical fiber probe system, which can measure the solids concentration and velocity at multi-points. The results reveal that with increasing U_g, local solids concentrations go through three evolution stages, reflecting a gradual regime transition process. Under all operating conditions, upflowing and descending particles co-exist at all measuring locations. The upflowing particle velocity is strong function of both superficial gas velocity and spatial position. However, the descending particle velocity mainly depends on superficial gas velocity. The bed radial symmetry and the effects of static bed height on the local flow structures are also investigated.     

AN EXPERIMENTAL STUDY OF LIQUID-PARTICLE FLOW IN CIRCULATING FLUIDIZED BED

A liquid-solid circulating fluidized bed (LSCFB) is operated at high liquid velocity, where particle entrainment is highly significant and between the conventional liquid fluidized bed and the dilute phase liquid transport regimes. LSCFB has potential applications in the fields of food processing, biochemical processing, and petrochemical and metallurgical processing. It is well known that the flow characteristics in a liquid-solid circulating fluidized bed are different from those of a conventional liquid-solid fluidized bed. The limited studies available in literature do not provide complete understanding of the flow structure in this typical regime.

In the present work, experiments were carried out in a 0.0762 m ID and 3 m height laboratory-scale liquid-solid circulating fluidized bed apparatus by using various solid particles and tap water as fluidizing medium. In the experimental setup, two distributors (specially designed) were used to monitor solid circulation rate in the riser. The effects of operating parameters, i.e., primary liquid flow rate in the riser (Up), solid circulation rate (G_s), and particle diameter (dp), were analyzed from the experimental data. Finally, a correlation was developed from the experimental data to estimate average solid holdup in the riser, and it was compared with present experimental and available data in the literature. They agree well with a maximum root-mean-square deviation of 7.83%.     

采用凝聚函数法对气固密相流化床流型转变过程的研究

本文对气固密相流化床中鼓泡流态化向湍动流态化转变的机理进行了研究.首次引入了凝聚函数分析法对密相压力脉动信号进行分析.结合时间域的统计分析,直接证明了这一流型转变乃是床内气泡的破碎程度超过合并程度所致.     

A consolidated flow regime map of upward gas fluidization

A new flow regime map, resulting from more fundamental studies on the hydrodynamics and new flow regimes, is proposed in response to more practical reclassifications of the existing regimes with the development of upward gas-solids fluidization systems. The previously reported flow regime maps and flow structures of some widely used fluidized beds are carefully examined. To better reflect the industrial applications, the fast fluidization regime is reclassified as high-density and low-density circulating fluidization regimes. A consolidated flow regime map is then proposed where the flow regimes of upward fluidization expand to include new types of fluidized beds such as circulating turbulent fluidized bed and high-density circulating fluidized bed. The proposed flow regime map consists of six flow regimes: bubbling, turbulent, circulating turbulent, high-density circulating and low-density circulating fluidization, and pneumatic transport. The transitions between the regimes are discussed with new correlations proposed for fluid catalytic cracking type particles. Analysis on the dominating phase in the different types of fluidized beds reveals the dynamic changeover from solids phase continuous in conventional low-velocity batch/“fixed” fluidization operations to gas phase continuous in high-velocity continuous/“moving” fluidization operations and provides more insights to the transitions between the flow regimes for industrial design and practice.