Packed beds mass transfer

Wilke and Hougan (1945) reported the mass transfer in beds of granular solids. Air was blown through a bed of porous celite pellets wetted with water, and by evaporating this water under adiabatic conditions, they reported gas-film coefficients for packed beds. In one run, the following data were reported ... 

Disadvantages of fixed beds include lower effectiveness factors because of the larger particle size and lower coefficients for gas-liquid and liquid-solid mass transfer. Fixed beds with two fluid phases are also difficult to scale up or scale down because of incomplete wetting and changes in gas and liquid flow distribution. Finally, there is a risk of temperature excursions with exothermic reactions in packed beds, since radial heat transfer is poor. Large reactors are often operated adiabatically, but hot spots may occur because of uneven flow distribution. 

Wagner I, Stichlmair J, Fair JR (1997) Mass transfer in beds of modem, high-efficiency random packings. Ind Eng Chem Res 36 227-237... 

From economical point of view die good packing must insure low capital investments and low operath] costs. The capital investments are connect with the intensity of the ma transfer processes in the packing, the operating costs - first of all with its pressiue drop. As a measure of the mass transfer intensity, the quantity of substance transferred in a imh of packing volume per imit of time ran be used. The possibility to increase this quantity is eai to te seen horn the basic equation of mass transfer in bed columns,... 

Fig. 20. Improved packing parameters ( ) for liquid mass transfer (a) ceramic Raschig rings (b) metal Raschig rings (c) ceramic Bed saddles (d) metal PaH...

External Fluid Film Resistance. A particle immersed ia a fluid is always surrounded by a laminar fluid film or boundary layer through which an adsorbiag or desorbiag molecule must diffuse. The thickness of this layer, and therefore the mass transfer resistance, depends on the hydrodynamic conditions. Mass transfer ia packed beds and other common contacting devices has been widely studied. The rate data are normally expressed ia terms of a simple linear rate expression of the form... 

N. Wakao, Heat and Mass Transfer in Packed Beds, Gordon Breach, New York, 1982. 

Mass Transfer and Useful Capacity. The term useful capacity, also referred to earlier as breakthrough capacity, differs from the equihbrium capacity shown on Figures 9a and 9b. The useful capacity is a measure of the total moisture taken up by a packed bed of adsorbent at the point where moisture begins to appear in the effluent. Thus the drying process cycle must be stopped before the adsorbent is fully saturated. The portion of the bed that is not saturated to an equihbrium level is called the mass-transfer 2one. 

To determine the mass-transfer rate, one needs the interfacial area in addition to the mass-transfer coefficient. For the simpler geometries, determining the interfacial area is straightforward. For packed beds of particles a, the interfacial area per volume can be estimated as shown in Table 5-27-A. For packed beds in distillation, absorption, and so on in Table 5-28, the interfacial area per volume is included with the mass-transfer coefficient in the correlations for HTU. For agitated liquid-liquid systems, the interfacial area can be estimated... 

Effects of Temperature on tiQ and tii The Stanton-number relationship for gas-phase mass transfer in packed beds,... 

A. Heat or mass transfer in packed bed for gases and liquids... 

Glaser and Thodos [Am. Jn.st. Chem. Eng. J., 4, 63 (1958)] give a correlation involving individual particle shape and bed porosity. Kunii and Suzuki [Jnt. ]. Heat Mass Tran.sfer, 10, 845 (1967)] discuss heat and mass transfer in packed beds of fine particles. 

End Effects Analysis of the mass-transfer efficiency of a packed cohimn should take into account that transfer which takes place outside the bed, i.e., at the ends of the packed sections. Inlet gas may veiy weU contact exit liquid below the bottom support plate, and exit gas can contact liquid from some types of distributors (e.g., spray nozzles). The bottom of the cohimn is the more hkely place for transfer, and SU-vey and KeUer [Chem. Eng. Prog., 62(1), 68 (1966)] found that the... 

Mass Transfer Mass transfer in plate and packed gas-liquid contactors has been covered earHer in this subsection. Attention nere will be limited to deep-bed contactors (bubble columns and agitated vessels). Theory underlying mass transfer between phases is discussed in Sec. 5 of this handbook. 

FIG. 16-9 General scheme of adsorbent particles in a packed bed showing the locations of mass transfer and dispersive mechanisms. Numerals correspond to mimhered paragraphs in the text 1, pore diffusion 2, solid diffusion 3, reaction kinetics at phase boundary 4, external mass transfer 5, fluid mixing. 

Rate equations are used to describe interphase mass transfer in batch systems, packed beds, and other contacting devices for sorptive processes and are formulated in terms of fundamental transport properties of adsorbent and adsorbate. 

FIG. 16"10 Sherwood mimher correlations for external mass-transfer coefficients in packed beds for e = 0.4 (adapted from Siiziild, gen. refs.). 

The linear driving force (LDF) approximation is obtained when the driving force is expressed as a concentration difference. It was originally developed to describe packed-bed dynamics under linear eqm-librium conditions [Glueckauf, Trans. Far Soc., 51, 1540 (1955)]. This form is exact for a nonlinear isotherm only when external mass transfer is controlling. However, it can also be used for nonlinear sys-... 

Axial Dispersion Effects In adsorption bed calculations, axial dispersion effects are typically accounted for by the axial diffusionhke term in the bed conservation equations [Eqs. (16-51) and (16-52)]. For nearly linear isotherms (0.5 < R < 1.5), the combined effects of axial dispersion and mass-transfer resistances on the adsorption behavior of packed beds can be expressed approximately in terms of an apparent rate coefficient for use with a fluid-phase driving force (column 1, Table 16-12) ... 

Gas/Liquid Mass Transfer This topic has been widely investigated for gas absorption in packed beds, usually countercurrent. One correlation for cocurrent flow in catalyst beds is by Sato et al. (First Pacific Chemical Engineering Congre.s.s, Pergamon, 1972, p. 187) ... 

Parameters a and b are related to the diffusion coefficient of solutes in the mobile phase, bed porosity, and mass transfer coefficients. They can be determined from the knowledge of two chromatograms obtained at different velocities. If H is unknown, b can be estimated as 3 to 5 times of the mean particle size, where a is highly dependent on the packing and solutes. Then, the parameters can be derived from a single analytical chromatogram. 

Superior surface utilization in mass and heat transfer, allowng shorter packed bed heights. Turn-down performance is superior over 2-in. and 3) -in. Pall rings. 

The three principal catalyst bed configurations are the pellet bed, the monolith, and the metallic wire meshes. An open structure with large openings is needed to fulfill the requirement of a low pressure drop even at the very high space velocities of 200,000 hr-1. On the other hand, packings with small diameters would provide more external surface area to fulfill the requirement for rapid mass transfer from the g .s stream to the solid surface. The compromise between these two ideals results in a rather narrow range of dimensions pellets are from to 1 in. in diameter, monoliths have 6 to 20 channels/in., and metallic meshes have diameters of about 0.004 to 0.03 in. 

The absorption of reactants (or desorption of products) in trickle-bed operation is a process step identical to that occurring in a packed-bed absorption process unaccompanied by chemical reaction in the liquid phase. The information on mass-transfer rates in such systems that is available in standard texts (N2, S6) is applicable to calculations regarding trickle beds. This information will not be reviewed in this paper, but it should be noted that it has been obtained almost exclusively for the more efficient types of packing material usually employed in absorption columns, such as rings, saddles, and spirals, and that there is an apparent lack of similar information for the particles of the shapes normally used in gas-liquid-particle operations, such as spheres and cylinders. 

By far the major portion of the available gas-absorption data have been obtained for countercurrent flow, which is the normal mode of operation for packed-bed absorbers. Special mention may be made of the results of Dodds et al. (D6), who examined mass transfer by the absorption of gas in liquid under cocurrent downward flow at flow rates higher than those corresponding to the flooding point for countercurrent operation. 

Zabor et al. (Zl) have described studies of the catalytic hydration of propylene under such conditions (temperature 279°C, pressure 3675 psig) that both liquid and vapor phases are present in the packed catalyst bed. Conversions are reported for cocurrent upflow and cocurrent downflow, it being assumed in that paper that the former mode corresponds to bubble flow and the latter to trickle-flow conditions. Trickle flow resulted in the higher conversions, and conversion was influenced by changes in bed height (for unchanged space velocity), in contrast to the case for bubble-flow operation. The differences are assumed to be effects of mass transfer or liquid distribution. 

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