Inlet conditions equilibrium conversion

Fig. 2.6. Conversion diagram with two exemplary trajectories for identical temperature and cell voltage, but with different inlet conditions. Both end up in the same attractor, which is the intersecting point of the reforming and the oxidation equilibrium lines.

This type of periodic operation allows for conversion improvement in reversible exothermic reactions [9], A cycle average inlet temperature for the conditions of continuous temperature oscillation can be substantially lower than the inlet temperature under steady-state conditions. This leads to a lower outlet temperature and higher equilibrium conversion for a reversible reaction. Better performance is achieved if temperature oscillations attenuate sufficiently during the passage through the catalyst bed [9]. 

A typical modem Claus sulphur recovery plant uses several reactors to achieve the equilibrium conversion of hydrogen sulfide. The complex gas nux-ture from the furnace is cooled to condense sulfur and then reheated before it enters the first catalyst reactor. There are generally three catalytic reactors in series contaitung a catalyst in series, with coolers at each reactor outlet to condense sulfur as it forms. Typical operating conditions are shown in Table 2.9. Inlet and outlet temperatures in each reactor are controlled at levels high enough to prevent condensation of sulfur on the catalyst. 

Fig. 10.17 Changes in CO conversion under Selox conditions as a function of temperature for two contact times (a) 170, (b) 48 ms. Thl thermodynamic equilibrium for Selox inlet composition (C0 02 H2 N2 = 1 2 20 77 vol%), Th2 thermodynamic equilibrium for similar Selox inlet composition but assuming that the water formed by CO/H2 oxidation is trapped by ceria.

CSBR). In this case the bioconversion is run under approximately steady-state conditions where the position of reaction equilibrium lies toward the products of the conversion. In this case the concentration of product (proportional to Sj — S0 ) at a given reactor residence time becomes a function of both the flow rate (Q) into the reactor and reactor volume, in addition to the factors discussed above for batch mode reactors (i.e., catalyst parameters and density, inlet substrate concentration S and outlet substrate concentration S0). 

The numerical case studied is derived from a flowsheet given in Stanford Research Institute Report 91, Isomerization of Paraffins for Gasoline. Since no kinetic information is given in this report, only reactor inlet and exit conditions, we will assume two different types of kinetics. In Case 1 we consider that the reaction is irreversible. An activation energy of 30,000 Btudb mol is used, and the preexponential factor is adjusted to give the same conversion reported in the SRI report. In Case 2 we assume that the reaction is reversible. The equilibrium constant decreases with increasing temperature because the reaction is exothermic. We also increase the size of the reactor so that the effluent leaves essentially at chemical equilibrium. 

As in the absence of reductant, NO2 concentration goes through a maximum when temperature increases, but this maximum is clearly below that observed without reductant. This can be explained if we consider that the thermodynamic equilibrium for NO oxidation should more appropriately be expressed with the reactor outlet temperature than to the reactor inlet temperature under adiabatic conditions the temperature increase is about 120 C for the combustion of 6000 ppmC hydrocarbon. This temperature shift is well suited to explain the NO oxidation curve in the presence of ethylene. It is less adapted in the presence of n-decane probably because decane oxidation is diffusion limited and reaches total conversion only at high temperature NO oxidation is not at equilibrium. 

Sulfuric acid plants are designed with optimized catalyst volumes and bed inlet temperatures to give a reasonable approach to equilibrium in each bed to achieve the maximum possible conversion of sulfur dioxide to sulfur trioxide. As shown by the examples in Table 2.8, this results in a significantly smaller volume in bed 1 than the remaining beds. The total catalyst volume used normally corresponds to a loading of 180-220 liters of catalyst per tonne of sulfuric acid produced per day although many plants use more, depending on conditions and the source of the sulfur dioxide. Lower volumes of catalyst are normally used in double-absorption units. 

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