Methane production reactor configurations

Partial oxidation of methane in the membrane reactor configuration shown in Figure 1 will not lead to higher yields of desired products than a plug flow reactor unless the diffusivity of the intermediate product, formaldehyde, is approximately four times that of methane. Presently available membranes that can withstand partial oxidation temperatures do not satisfy this criterion. 

Figure 1. Absolute production rate of HCN vs. feed rate of methane (both normalized by dividing by G, the most common flow rate of nitrogen. G = 0.0171 gram moles/sec. = 0.383 liter (STP)/sec.) at three net power levels 3522 watts—O 6992 watts— and 9782 watts— . Old data, reactor configuration of Figure 5, (0 inch 2 inch)...

The studies have proven that both MRs can realize better performance in terms of methane conversion and hydrogen production than the CRs, working also at milder operating conditions. By making a comparison between the two reactor configurations, it has been shown that a PBMR has a very simple configuration whereas an FBMR is typically characterized by enhanced mass and heat transfer rates, which favor more uniform temperature profiles. Nevertheless, possible bubble-to-emulsion mass transfer... 

Porous alumina tube externally coated with a MgO/PbO dense film (in double pipe configuration), tube thickness 2.5 mm, outer diameter 4 mm, mean pore diameter 50 nm, active film-coated length 30 mm. Feed enters the reactor at shell side, oxygen at tube side. Oxidative methane coupling, PbO/MgO catalyst in thin film form (see previous column). r-750X,Pr ed 1 bar. Conversion of methane <2%. Selectivity to Cj products > 97%. Omata et al. (1989). The methane conversion is not given. Reported results are calculated from permeability data. 

To circumvent such problems, up to two-thirds of the acid gas has been removed from the flame and sent directly to the catalytic reactor, the so-called "long bypass" scheme. This strategy has worked with mixed success in natural gas plants, where hydrocarbons in the acid gas are C1-C3 paraffins, chiefly methane, and no olefins. Such a configuration has been put forward for use in coal gasification plants. In the writer s opinion it is unworkable if the bypassed acid gas contains even traces of olefins or aromatics these compounds react with SO2 to form tarry products which foul the catalyst and discolor the product sulfur. 

A glass MSR was used to perform the dehydration of ethanol. The microchan-nel of size 200 X 80 pm deep X 30 mm (in a Z shaped configuration) was produced by photohthographic etching [71]. A sulfated zirconia catalyst immobihzed over the surface of the top cover block. In addition, a NiCr wire was immobilized in the reactor cover as a heating device. At a reaction temperature of 155 °C and a flow rate of 3 plmin the main products were 68% ethylene, 16% ethane, and 15% methane. A further increase of the residence time resulted in a reaction progress beyond dehydration to almost complete cracking of the ethanol to methane. 

In the catalytic reformer, the reactants (CH4, H2O, and O2) are mixed and then enter directly to the catalyst bed. Carbon formation is avoided with this process, because there is not enough residence time in the empty space above the catalyst bed, even when the methane is preheated at low temperatures. The necessary amount of steam and oxygen is lower, and as a result the ratio between CO and CO2 in the product is high. The drawbacks of this configuration are the high mechanical and thermal loads on the catalyst close to the entrance of the reactor. Moreover, catalysts disintegration may occur due to the temperature fluctuation during the start-up and shut-down of process. 

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