Catalytic external mass transfer

A final, obvious but important, caution about catalyst film preparation Its thickness and surface area Ac must be low enough, so that the catalytic reaction under study is not subject to external or internal mass transfer limitations within the desired operating temperature range. Direct impingement of the reactant stream on the catalyst surface1,19 is advisable in order to diminish the external mass transfer resistance. 

Traditionally, an average Sherwood number has been determined for different catalytic fixed-bed reactors assuming constant concentration or constant flux on the catalyst surface. In reality, the boundary condition on the surface has neither a constant concentration nor a constant flux. In addition, the Sh-number will vary locally around the catalyst particles and in time since mass transfer depends on both flow and concentration boundary layers. When external mass transfer becomes important at a high reaction rate, the concentration on the particle surface varies and affects both the reaction rate and selectivity, and consequently, the traditional models fail to predict this outcome. 

In catalytic gas-liquid-solid systems mass transfer is more complex. The catalyst particles are present in the liquid phase. The expression for the rate of mass transfer from the gas to the liquid is identical to that for systems without a solid catalyst (Eqn. 5.4-67). However, now also mass transfer from the liquid to the solid surface (external mass transfer) and inside the particle (internal mass transfer) have to be considered. 

Before terminating the discussion of external mass transfer limitations on catalytic reaction rates, we should note that in the regime where external mass transfer processes limit the reaction rate, the apparent activation energy of the reaction will be quite different from the intrinsic activation energy of the catalytic reaction. In the limit of complete external mass transfer control, the apparent activation energy of the reaction becomes equal to that of the mass transfer coefficient, typically a kilocalorie or so per gram mole. This decrease in activation energy is obviously... 

Packed-bed reactors, 21 333, 352, 354 Packed beds, 25 718 Packed catalytic tubular reactor design with external mass transfer resistance, 25 293-298 nonideal, 25 295... 

There are a number of examples of tube waU reactors, the most important being the automotive catalytic converter (ACC), which was described in the previous section. These reactors are made by coating an extruded ceramic monolith with noble metals supported on a thin wash coat of y-alumina. This reactor is used to oxidize hydrocarbons and CO to CO2 and H2O and also reduce NO to N2. The rates of these reactions are very fast after warmup, and the effectiveness factor within the porous wash coat is therefore very smaU. The reactions are also eternal mass transfer limited within the monohth after warmup. We wUl consider three limiting cases of this reactor, surface reaction limiting, external mass transfer limiting, and wash coat diffusion limiting. In each case we wiU assume a first-order irreversible reaction. 

When a biocatalyst is immobilized on or within a solid matrix, mass transfer effects may exist because the substrate must diffuse from the bulk solution to the immobilized biocatalyst. If the biocatalyst is attached to non-porous supports there are only external mass transfer effects on the catalytically active outer surface in the reaction solution, the supports are surrounded by a stagnant film and substrate and product are transported across this Nemst layer by diffusion. The driving force for this diffusion is the concentration difference between the surface and the bulk concentration of substrate and product. 

At catalytically active centers in the center of carrier particles, external mass transfer (film diffusion) and/or internal mass transfer (pore diffusion) can alter or even dominate the observed reaction rate. External mass transfer limitations occur if the rate of diffusive transport of relevant solutes through the stagnating layer at a macroscopic surface becomes rate-limiting. Internal mass transfer limitations in porous carriers indicate that transport of solutes from the surface of the particle towards the active site in the interior is the slowest step. 

In order to avoid the unfavorable process conditions, different flue-gas treatment processes for combustion plants based on catalytic filters were developed, which combine fly-ash removal with SCR of ISKh with NH3 [4—8], The advantages of these processes are space and treatment-cost savings, reduced internal and external mass transfer resistances compared to honeycomb SCR catalysts, heat recovery from offgases with good efficiency, and low corrosion problems due to the removal of both dust and NOx at high temperatures. 

Cybulski and Moulijn [27] proposed an experimental method for simultaneous determination of kinetic parameters and mass transfer coefficients in washcoated square channels. The model parameters are estimated by nonlinear regression, where the objective function is calculated by numerical solution of balance equations. However, the method is applicable only if the structure of the mathematical model has been identified (e.g., based on literature data) and the model parameters to be estimated are not too numerous. Otherwise the estimates might have a limited physical meaning. The method was tested for the catalytic oxidation of CO. The estimate of effective diffusivity falls into the range that is typical for the washcoat material (y-alumina) and reacting species. The Sherwood number estimated was in between those theoretically predicted for square and circular ducts, and this clearly indicates the influence of rounding the comers on the external mass transfer. 

The differences between the TBR and the MR originate from the differences in catalyst geometry, which affect catalyst load, internal and external mass transfer resistance, contact areas, as well as pressure drop. These effects have been analyzed by Edvinsson and Cybulski [ 14,26] via computer simulations based on relatively simple mathematical models of the MR and TBR. They considered catalytic consecutive hydrogenation reactions carried out in a plug-flow reactor with cocurrent downflow of both phases, operated isothermally in a pseudo-steady state all fluctuations were modeled by a corresponding time average ... 

Owing to the comparatively small size of the pores (up to 100 p.m, compared to a pitch of a few millimeters for the honeycomb channels) and the small thickness of the catalyst layer (a few microns, compared to some tenths of a millimeter for the catalytic wall of the honeycomb channels), both internal and external mass transfer limitations to NO conversion in catalytic filters can easily be neglected. An efficiency factor equal to unity can thus be assumed with confidence for NO reduction, contrary to honeycomb catalysts, for which this parameter is hardly higher than 0.05% at the conventional operating temperatures (320-380 C). 

It may seem reasonable to employ 1-hexene Itself as the solvent medium to allow higher operating temperatures. The critical temperatures of 1-hexene and Its Isomers are relatively high (231 -248 C) thus favoring reaction kinetics. However, the overall rate of the catalytic Isomerization of pure 1-hexene may become limited by external mass-transfer and Internal pore-dlffusion resistances at the higher operating temperatures (T - 1.1-1.2 T ). 

H has the dimensions of a length and has been called the height of a reactor unit, by analogy with heights of transfer units and equivalent theoretical plates. Interpret Eq. (6.5.4) to show that H is the sum of a height for external mass transfer HTU) and a term dependent on the reaction, the so-called height of a catalytic unit (HCU). Examine the contribution of these terms when the mass transfer, diffusion, and kinetic regimes are dominant. 

Consider a case where the true order of the surface reaction is 2 [according to Eq. (10-2)] but the rate is diffusion controlled, so that Eq. (10-3) is applicable. Experimental data plotted as rate vs would yield a straight line. If diffusion were not considered, and Eq. (10-2) were used to interpret the data, the order would be identified as unity—a false conclusion. This simple example illustrates how erroneous conclusions can be reached about kinetics of a catalytic reaction if external mass transfer is neglected, a... 

Often the global reaction rate of heterogeneous catalytic reactions is affected by the diffusion in the pore and the external mass-transfer rate of the reactants and the products. When the diffusion in the pores is not fast, a reactant concentration profile develops in the interior of the particle, resulting in a different reaction rate at different radial locations inside the catalytic pelet. To relate the global reaction rate to various concentration profiles that may develop, a kinetic effectiveness factor is defined [1, 3,4,7, 8] by... 

External mass transfer In general, the thickness of the catalyst layer will be kept sufficiently small to avoid the influence of internal mass transfer on the kinetics. In this case, only the transfer of the reactants from the bulk of the fluid to the catalytic wall must be considered. The radial velocity profile in a single channel develops from the entrance to the position where a complete Poiseuille profile is established (provided that the flow is laminar). The length of the entrance zone depends on the Reynolds number and can be estimated from the following empirical relationship [85,86] ... 

In three-phase reactors, one of the main problems is often the mass transport limitations, which may reflect internal as well as external mass transfer resistances. The use of filamentous catalytic materials for multiphase reactions may help reduce or even avoid mass transfer limitations [63,132,133]. Filamentous woven cloths made of glass, composite mixed oxides, metallic alloys, or activated carbon (Figure 18) can be used as supports for active components such as platinum, palladium, or transition metal oxides. The diameters of the filaments are of the order of several micrometers and correspond to the typical diameters of catalysts that are suspended in the reaction medium. By using such small diameters, internal mass transfer limitations can be avoided. 

Gas-solid (catalytic) reactions. Mass transfer is likely to be more important within the pellet than in the external film, and heat transfer more important in the film than within the pellet. In other words, intraphase mass transfer and interphase heat transfer would normally be the dominant transport processes. Thus the pellet can reasonably be assumed to be isothermal. 

Of course most industrial catalytic reactors (except fluidized bed catalytic reactors) use relatively large particle sizes for the catalyst to avoid the excessive pressure drop associated with fine particles. This gives rise to intraparticle mass transfer resistances. However, in most industrial reactors, but not all, the gas flowrate is quite large rendering external mass transfer resistances usually negligible. 

The transfer of molecules from the bulk phase to the location where reaction occurs (catalytically or thermally) may significantly influence the overall reaction rate. When the transfer from the bulk fluid phase to the catalyst is limiting, this is referred to as external mass transfer limitation. If diffusion of reactants or products in the pores of the catalyst is slow, it is termed internal diffusion limitation (Fig. 3.1). Both of these effects commonly occur under relevant conditions. 

Step 1. Reactants enter a packed catalytic tubular reactor, and they must diffuse from the bulk fluid phase to the external surface of the solid catalyst. If external mass transfer limitations provide the dominant resistance in this sequence of diffusion, adsorption, and chemical reaction, then diffusion from the bulk fluid phase to the external surface of the catalyst is the slowest step in the overall process. Since rates of interphase mass transfer are expressed as a product of a mass transfer coefficient and a concentration driving force, the apparent rate at which reactants are converted to products follows a first-order process even though the true kinetics may not be described by a first-order rate expression. Hence, diffusion acts as an intruder and falsifies the true kinetics. The chemical kineticist seeks to minimize external and internal diffusional limitations in catalytic pellets and to extract kinetic information that is not camouflaged by rates of mass transfer. The reactor design engineer must identify the rate-limiting step that governs the reactant product conversion rate. 

If integration is performed along the catalytically active perimeter in one quadrant only (i.e., the xy plane where both x and y are positive), then file complete circumference of the tube on the right side of the preceding equation is replaced by tzR/2, and S = nR /4. However, the final result is unchanged. See Problem 30-7 for a continuation of this analysis and a solution of the quasi-macroscopic plug-flow mass balance in the presence of significant external mass transfer resistence. 

This suggests that external mass transfer resistance is not very important for (1) low rates of reactant conversion, (2) small catalytic pellets with large external surface-to-volume ratios, and (3) high gas-phase flow rates through the packed... 

The important dimensionless parameter that determines the significance of external mass transfer resistance for nth-order irreversible chemical kinetics in packed catalytic tubular reactors was introduced in equation (30-63) as a = iS(CA.iniet)" Simple algebraic manipulation allows one to relate a to the interpellet Damkohler number, the effectiveness factor, the mass transfer Peclet number, and a few other dimensionless parameters. For example, let the coefficient of the chemical reaction term in the dimensionless mass transfer equation be defined as follows ... 

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