Gas-particle mass transfer

It must be emphasised that particle mixing remains an essential consideration in the use of fluidization in other operations such as drying and granulation. The mixing in a fluidized bed of particles of 

Convection describes the movement of groups of particles from one place to another within the mixer volume because of the direct action of an impeller or a moving device within the mixer body. As in convection within fluids, this is likely to be a more significant effect than diffusion but diffusional effects will still be present. 

The shear mechanism operates when slipping planes are formed within the particulate mass, perhaps because of the action of a blade, which in turn allow particles to exploit new void spaces through which particles can then diffuse. 

The dispersion model of solids mixing (Kunii and Levenspiel, 1991) is based on Tick s equafion which describes unsfeady-sfafe mass transfer in one dimension. Thus 

Here the dispersion coefficient (m s ) is directly proportional to the superficial gas velocity, over the range 0.1-1 ms h For larger diameter beds, in the range 0.1-3 m (and gas velocities between 0.2 and 0.5 ms ), they suggested 

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The experimental mass transfer coefficients for gas-particle mass transfer in a fluidized bed (Figure 2.2), as summarised by Kunii and... 

Figure 2.2 Experimental gas-particle mass transfer coefficients. Adapted from Kunii, D. and Levenspiel, O., Fluidization engineering, 1991, with permission from Elsevier.

Table 2.2 Example calculation of gas-particle mass transfer coefficients.

FIG. 7-13 Concentration profiles with reaction control ( = 1, in absence of gas particle mass-transfer resistance. [From Wen, oncat-alytic Heterogeneous Solid-Fluid Reaction Models, Ind. Eng. Chem. 60(9) 34-54 (1968), Fig. 11.]... 

Table 4.17 Expressions for gas-particle mass transfer n , g, Mq and Haq molecule density of the same substance far from the particle, close to the particle, at particle surface and within the particle (droplet) p — gas partial pressure far from the droplet, c g - aqueous-phase concentration, k - mass transfer coefficient (recalculable into spjecific rate constant) g - gas-phase, aq - aqueous-phase, het - interfadal layer (chemistry), in - interfacial layer (transport), coll - collision, ads — adsorption (surface striking), sol - dissolution, diff -diffusion in gas-phase, des - desorption.

Each stage of particle formation is controlled variously by the type of reactor, i.e. gas-liquid contacting apparatus. Gas-liquid mass transfer phenomena determine the level of solute supersaturation and its spatial distribution in the liquid phase the counterpart role in liquid-liquid reaction systems may be played by micromixing phenomena. The agglomeration and subsequent ageing processes are likely to be affected by the flow dynamics such as motion of the suspension of solids and the fluid shear stress distribution. Thus, the choice of reactor is of substantial importance for the tailoring of product quality as well as for production efficiency. 

Several reported chemical systems of gas-liquid precipitation are first reviewed from the viewpoints of both experimental study and industrial application. The characteristic feature of gas-liquid mass transfer in terms of its effects on the crystallization process is then discussed theoretically together with a summary of experimental results. The secondary processes of particle agglomeration and disruption are then modelled and discussed in respect of the effect of reactor fluid dynamics. Finally, different types of gas-liquid contacting reactor and their respective design considerations are overviewed for application to controlled precipitate particle formation. 

Sada, E., Kumazawa, H., Lee, C. and Fujiwara, N., 1985. Gas-liquid mass transfer characteristics in a bubble column with suspended sparingly soluble fine particles. Industrial and Engineering Chemistry Process Design and Development, 24, 255-261. 

Nore, O., Briens, C., Margaritis, A., and Wild, G., Hydrodynamics, Gas-Liquid Mass Transfer and Particle-Liquid Heat and Mass Transfer in a Three-Phase Fluidized Bed for Biochemical Process Applications, Chem. Eng. Sci., 47 3573 (1992)... 

An important difference between a shrinking particle reacting to form only gaseous product(s) and a constant-size particle reacting so that a product layer surrounds a shrinking core is that, in the former case, there is no product or ash layer, and hence no ash-layer diffusion resistance for A. Thus, only two rate processes, gas-film mass transfer of A, and reaction of A and B, need to be taken into account. 

Assume that the particle is spherical and isothermal, that both gas-film mass transfer resistance and reaction resistance are significant, and that the Ranz-Marshall correlation for k g is applicable. Do not make an assumption about particle size, but assume the reaction is first-order. 

The performance of a reactor for a gas-solid reaction (A(g) + bB(s) -> products) is to be analyzed based on the following model solids in BMF, uniform gas composition, and no overhead loss of solid as a result of entrainment. Calculate the fractional conversion of B (fB) based on the following information and assumptions T = 800 K, pA = 2 bar the particles are cylindrical with a radius of 0.5 mm from a batch-reactor study, the time for 100% conversion of 2-mm particles is 40 min at 600 K and pA = 1 bar. Compare results for /b assuming (a) gas-film (mass-transfer) control (b) surface-reaction control and (c) ash-layer diffusion control. The solid flow rate is 1000 kg min-1, and the solid holdup (WB) in the reactor is 20,000 kg. Assume also that the SCM is valid, and the surface reaction is first-order with respect to A. 

As shown in Example 22-3, for solid particles of the same size in BMF, the form of the reactor model resulting from equation 22.2-13 depends on the kinetics model used for a single particle. For the SCM, this, in turn, depends on particle shape and the relative magnitudes of gas-film mass transfer resistance, ash-layer diffusion resistance and surface reaction rate. In some cases, as illustrated for cylindrical particles in Example 22-3(a) and (b), the reactor model can be expressed in explicit analytical form additional results are given for spherical particles by Levenspiel(1972, pp. 384-5). In other f l cases, it is convenient or even necessary, as in Example 22-3(c), to use a numerical pro-... 

Ka can be defined as a gas-phase transfer coefficient, independent of the liquid layer, when the boundary concentration of the gas is fixed and independent of the average gas-phase concentration. In this case, the average and local gas-phase mass-transfer coefficients for such gases as sulfur dioxide, nitrogen dioxide, and ozone can be estimated from theoretical and experimental data for deposition of diffusion-range particles. This is done by extending the theory of particle diffusion in a boundary layer to the case in which the dimensionless Schmidt number, v/D, approaches 1 v is the kinematic viscosity of the gas, and D is the molecular diffusivity of the pollutant). Bell s results in a tubular bifurcation model predict that the transfer coefficient depends directly on the... 

Another type of mass transfer equipment, shown in Figure 6.2d, is normally referred to as the packed- (fixed-) bed. Unlike the packed column for gas-liquid mass transfer, the packed-bed column is used for mass transfer between the surface of packed solid particles (e.g., catalyst particles or immobilized enzyme particles) and a single-phase liquid or gas. This type of equipment, which is widely used as reactors, adsorption columns, chromatography columns, and so on, is discussed in greater detail in Chapters 7 and 11. 

The lack of gas-side mass-transfer data at elevated pressure for countercurrent fixed beds packed with large and small solid particles should be underlined. 

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