The Catalytic Wall Reactor

This problem may look at first to be the same geometry as for diffusion in a single pore, but this situation is quite drSerent hr the single pore we had reaction controlled by diffusion down the pore, while in the tube wall reactor we have convection of reactants down the tube. 

Here (A/ V) is the waU area per unit volume of the tube. For a cylindrical tube the area per 

Note the different ds in these problems. While d always signifies diameter, D is reactor diameter (or diameter of a cylinder or sphere because that notation is used in mass transfer rather than the radius R), dp is the diameter of a single pore, and d is the diameter of a single tube in a tube waU reactor. 

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. 

If the process is reaction limiting, the composition is given directly by the preceding 

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It is evident, however, that this problem can be much more comphcated than either the wetted wall column or the catalytic wall reactor, because it combines the complexities of both. In fact, there are numerous additional complexities with this reactor beyond those simplified cases. 

As with the falling film reactor, the rate of mass transfer to the catalyst goes as R, while the size of the reactor goes as R, so this reactor becomes very inefficient except for very small-diameter tubes. However, we can overcome this problem, not by using many small tubes in parallel, but by allowing the gas and liquid to flow (trickle) over porous catalyst pellets in a trickle bed reactor rather than down a vertical wall, as in the catalytic wall reactor. 

To compare the catalytic wall reactor with a packed bed, correct criteria must be chosen [18]. For both the reactors, the outer catalyst surface per void (V oy) volume and the space-time must be identical. Under these conditions, the following relationship between the diameter of the microchannel and the particle diameter holds... 

The catalytic wall reactor with channel diameter in the range of 50-1000 pm and a length dependent on the reaction time required circumvents the shortcomings of micro packed beds. This is discussed in more detail in Section 6.5.4. However, in most of the cases, the catalytic surface area provided by the walls alone is insufficient for the chemical transformation and, therefore, the SSA has to be increased by the chemical treatment of the channel walls, or by coating them with highly porous support layers. The thickness of the layer 5 3, depends on catalytic activity. In general, the layer thickness is sufficiently small to avoid internal heat and mass transfer influences. Catalytic layers can be obtained by using a... 

Munder, Rihko-Stmckmann, and Sundmacher (2007) investigated the performance of a bilayer-solid electrolyte membrane reactor in steady-state and forced-periodic operation modes. The obtained results showed that the solid electrolyte membrane reactor has a slightly higher yield under optimal operation conditions compared with the catalytic wall reactor, and forced-periodic operation of the bilayer-solid electrolyte... 

Unlike SRE, the POE reaction for H2 production has been reported so far only by a few research groups.101104-108 While Wang et al. os and Mattos et r//.104-106 have studied the partial oxidation of ethanol to H2 and C02 (eqn (18)) at lower temperatures, between 300 and 400 °C using an 02/EtOH molar ratio up to 2, Wanat et al.101 have focused on the production of syngas (eqn (19)) over Rh/Ce02-monolith catalyst in a catalytic wall reactor in millisecond contact time at 800 °C. Depending on the nature of metal catalyst used and the reaction operating conditions employed, undesirable byproducts such as CH4, acetaldehyde, acetic acid, etc. have been observed. References known for the partial oxidation of ethanol in the open literature are summarized in Table 6. 

Figure 12-13 A falling film catalytic wall reactor in which reactant in the gas must diffuse through a liquid film to react L...

This intermediate scale affords a preliminary validation of the intrinsic kinetics determined on the basis of microreactor runs. For this purpose, the rate expressions must be incorporated into a transient two-phase mathematical model of monolith reactors, such as those described in Section III. In case a 2D (1D+ ID) model is adopted, predictive account is possible in principle also for internal diffusion of the reacting species within the porous washcoat or the catalytic walls of the honeycomb matrix. 

Kolios et al. [106] performed an extensive study revealing that coupling of exothermic and endothermic reactions is possible under safe and stable operation conditions only in catalytic wall reactors and not in coupled packed beds. In the latter case instability and thermal runaway of the reactor may occur. Additionally, the two... 

Fig. 1. The multiple scales in the catalytic monolith reactor (a) catalytic monolith (10 cm), (b) channel with catalyst washcoat on the walls (1mm), (c) SEM image of the washcoat layer (10 pm), (d) TEM image of meso-porous y-Al203 with dispersed Pt (200 nm).

The mass transfer equation is written in terms of the usual assumptions. However, it must be considered that because the concentration of the more abundant species in the flowing gas mixture (air), as well as its temperature, are constant, all the physical properties may be considered constant. The only species that changes its concentration along the reactor in measurable values is PCE. Therefore, the radial diffusion can be calculated as that of PCE in a more concentrated component, the air. This will be the governing mass transfer mechanism of PCE from the bulk of the gas stream to the catalytic boundaries and of the reaction products in the opposite direction. Since the concentrations of nitrogen and oxygen are in large excess they will not be subjected to mass transfer limitations. The reaction is assumed to occur at the catalytic wall with no contributions from the bulk of the system. Then the mass balance at any point of the reactor is... 

A major problem in using microstructured reactors for heterogeneously catalyzed gas-phase reactions is how to introduce the catalytic active phase. The possibilities are to (i) introduce the solid catalyst in the form of a micro-sized packed bed, (ii) use a catalytic wall reactor or (iii) to use novel designs. Kiwi-Minsker and Renken [160] have discussed in detail these alternatives. 

H. Redlingshofer, O. Krocher, W. Bock, K. Huthmacher, G. Emig, Catalytic wall reactor as a tool for isothermal investigations in the heterogeneously catalyzed oxidation of propene to acrolein, Ind. Eng. Chem. Res. 41 (2002) 1445. 

This model does not apply to a porous wall tubular reactor. In this case, we must account for solution losses along the tube length, and a radial convective term must be included. Again, the enzymatic reaction at the catalytic wall enters into the model as a boundary condition. 

A simple estimation of the temperature profile inside a tube reactor starts from an energy balance for the system [7]. Since we are particularly interested in the performances of a tube reactor and a catalytic wall reactor the following simplifying assumptions are made ... 

The use of microstructured catalytic wall reactors offers an interesting option for the revamping of existing plants. The key idea of this so-called booster concept... 

As an example of the decreasing efficiency caused by external and internal mass transfers, we consider an irreversible first-order reaction and a porous catalyst layer. The situation corresponds to a catalytic wall reactor. The relative importance between external and internal mass transfers is characterized by the ratio of the diffusion time in the porous layer tp and the characteristic time for external mass transfer called the Biot number, Bi = t /t - (L /DJk a for mass transfer. 

If the mass transfer is accompanied by a chemical reaction at the catalyst surface on the reactor wall, the mass transfer depends on the reaction kinetics [55]. For a zero-order reaction, the rate is independent of the concentration and the mass flow from the bulk to the wall is constant, whereas the reactant concentration at the catalytic wall varies along the reactor length. For this situation the asymptotic Sh in circular tube reactors becomes Sh. = 4.36 [55]. The same value is obtained when reaction rates are low compared to the rate of mass transfer. If the reaction rate is high (very fast reactions), the concentration at the reactor wall can be approximated to zero within the whole reactor and the asymptotic value for Sh is = 3.66. As a consequence, the Sh in the reacting system depends on the ratio of the reaction rate to the rate of mass transfer characterized by the second Damkohler number defined in Equation 6.11. 

Nevertheless, there are several constraints hampering the use of microstruc-tured devices for fluid-solid reactions. In the catalytic reactions, the performance is very adversely affected by catalyst deactivation. Effective in situ catalyst regeneration thus becomes necessary, as the simple catalyst change practiced in conventional reactors is usually no longer an option. The thickness of the catalytic wall is often greater than the internal diameter of the channel and, therefore, may impede heat transfer for highly exothermic reactions leading to nonisothermal behavior. 

The gas-liquid-sohd reactions are carried out in various types of reactors, such as packed beds, fluidized/slurry, and catalytic wall reactors (Figure 8.1). The advantages and limitations of these reactors are described in Table 8.1. Compared to fluid-sohd systems, an additional phase makes it difficult to predict flow patterns... 

A purchasable cross-flow heat exchanger for application in laboratory-, pilot- and production-scale plants was developed by FZK. By incorporation of a catalyst on the quadratic plates inside the heat exchanger, it can also be used as a catalytic wall reactor. Operating conditions up to 850 °C (stainless steel) and pressures of more than 100 bar are possible, and the specific inner surface area is up to 30 000 m m. The reactors can be obtained in many materials and three different sizes with a maximum flow of 6500kgh (water). Therefore, the reactors can be adjusted for various processes, and all types of catalyst deposition techniques are possible [111]. This reactor has already been applied to the catalytic oxidation of H2 by Janicke et al. [112], for example. 

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