Methanol, reactor simulation

Baddour et al. [26] in their simulation of the TVA ammonia-synthesis converter, already discussed in Sec. 11.5.e, found that in steady-state operation the temperature difference between the gas and the solid at the top, where the rate of reaction is a maximum, amounts to only 2.3°C and decreases as the gas proceeds down the reactor to a value of 0.4°C at the outlet. In the methanol reactor simulated in Sec. 11.9.b the difference between gas and solid temperature is of the order of 1 C. This may not be so with highly exothermic and fast reactions involving a component of the catalyst as encountered in the reoxidation of Fe and Ni catalysts used in ammonia synthesis and steam reforming plants or involving material deposited on the catalyst, coke for example. 

In batch mode, conversions of up to 92.5% were achieved under the above mentioned conditions (Table 6.5). More complete conversions for a continuous reactor design would require intermittent removal of excess methanol and water, possible by flashing the wet methanol. To simulate a continuous process with removal of water a batch experiment was subjected to three repeat cycles of flashing water and methanol, resulting in an acid value 0.5 mg KOH/g for the crude ester product. 

The operating conditions used for the methanol Bench Scale Reactor simulations are listed in the tables 11.3-11.5. 

For Example 5.7, use a simulator to graph the effluent temperature of the methanol reactor as a function of the H2/CO ratio. 

Bos ANR, Tromp PJJ, Akse HN Conversion of methanol to lower olefins. Kinetic modeling, reactor simulation, and selection, Ind Eng Chem Res 34 3808—3816, 1995. 

This program helps calculate the rate of methanol formation in mol/m s at any specified temperature, and at different hydrogen, carbon monoxide and methanol concentrations. This simulates the working of a perfectly mixed CSTR specified at discharge condition, which is the same as these conditions are inside the reactor at steady-state operation. Corresponding feed compositions and volumetric rates can be calculated from simple material balances. 

Also a simulation of the flow field in the methanol-reforming reactor of Figure 2.21 by means of the finite-volume method shows that recirculation zones are formed in the flow distribution chamber (see Figure 2.22). One of the goals of the work focused on the development of a micro reformer was to design the flow manifold in such a way that the volume flows in the different reaction channels are approximately the same [113]. In spite of the recirculation zones found, for the chosen design a flow variation of about 2% between different channels was predicted from the CFD simulations. In the application under study a washcoat cata-... 

A reformate flow rate of 25-175 Ndm3 min-1, simulating methanol steam reformer product gases, and 2.5-17.5 Ndm3 min-1 air were fed to the reactors. The simulated reformate was composed of 68.9% H2, 0.6% CO, 22.4% C02, 6.9% H20 and 0.4% CH3OH, the last to simulate incomplete conversion. The carbon monoxide output of the single reactors and of both switched in series is shown in Figure 2.72. The CO output of the two reactors switched in series was <10 ppm and the optimum air volume split between the first and second reactors was determined as 70/30. 

For reformate flow rates up to 400 Ndm3 min-1, the CO output was determined as < 12 ppm for simulated methanol. The reactors were operated at full load (20 kW equivalent power output) for -100 h without deactivation. In connection with the 20 kW methanol reformer, the CO output of the two final reactors was < 10 ppm for more than 2 h at a feed concentration of 1.6% carbon monoxide. Because the reformer was realized as a combination of steam reformer and catalytic burner in the plate and fin design as well, this may be regarded as an impressive demonstration of the capabilities of the integrated heat exchanger design for fuel processors in the kilowatt range. 

To conclude our examples of Aspen Dynamics simulation of tubular reactor systems, we study a very important industrial process for the production of methanol from synthesis... 

M. C. Bjorklund and R. W. Carr, Enhanced methanol yields from the direct partial oxidation of methane in a simulated countercurrent moving bed chromatographic reactor. Indust. Engng. Chem. 

Example 5.15 Retrofits of distillation columns by thermodynamic analysis The synthesis of methanol takes place in a tube reactor in section 3 in the methanol plant shown in Figure 5.7. The reactor outlet is flashed at 45°C and 75 bar, and the liquid product (stream 407) containing 73.45 mol% of methanol is fed into the separation section (see Figure 5.8), where the methanol is purified. Stream 407 and the makeup water are the feed streams to the section. Table 5.2 shows the properties and compositions of the streams in section 3. The converged simulations are obtained from the Redlich-Kwong-Soave method to estimate the vapor properties, while the activity coefficient... 

Cybulski et al. [39] have studied the performance of a commercial-scale monolith reactor for liquid-phase methanol synthesis by computer simulations. The authors developed a mathematical model of the monolith reactor and investigated the influence of several design parameters for the actual process. Optimal process conditions were derived for the three-phase methanol synthesis. The optimum catalyst thickness for the monolith was found to be of the same order as the particle size for negligible intraparticle diffusion (50-75 p.m). Recirculation of the solvent with decompression was shown to result in higher CO conversion. It was concluded that the performance of a monolith reactor is fully commensurable with slurry columns, autoclaves, and trickle-bed reactors. 

Other steps used in the model assume that the heterogeneous conversion of methane is limited to the gas-phase availability of oxygen, O2 adsorption is fast relative to the rate of methane conversion, and heat and mass transports are fast relative to the reaction rates. Calculations for the above model were conducted for a batch reactor using some kinetic parameters available for the oxidative coupling of methane over sodium-promoted CaO. The results of the computer simulation performed for methane dimerization at 800 °C can be found in Figure 7. It is seen that the major products of the reaction are ethane, ethylene, and CO. The formation of methanol and formaldehyde decreases as the contact time increases. 

TABLE 11.2 Methanol Synthesis Reactor Comparison of simulations and experimental data (from Wu and Gidaspow, 2000)... 

Wu, Y. and Gidaspow, D. (2000), Hydrodynamic simulation of methanol synthesis in gas-liquid slurry bubble column reactors, Chem. Eng. Sci., 55, 573-587. 

The results from a simulation of the methanol process with external cooling (non-adiabatic process) are given in Fig 11.4. The predicted profiles show that methanol is produced at the expense of CO, CO2 and H2. The temperature gradient at the reactor entrance is very steep. The temperature increases to about 550 K in the center of the tube, but near the walls the maximum... 

The kinetic theory model was extended to include the effect of the mass transfer coefficient between the liquid and the gas and the water gas shift reaction in the slurry bubble column reactor. The computed granular temperature was around 30 cm /sec and the computed catalyst viscosity was closed to 1.0 cp. The volumetric mass transfer coefficient estimated by the simulation has a good agreement with experimental values shown in the literature. The optimum particle size was determined for maximum methanol production in a SBCR. The size was about 60 - 70 microns, found for maximum granular temperature. This particle size is similar to FCC particle used in petroleum refining. 

The first methanol was vaporised and reacted to produce hydrocarbons in the first conversion reactor as anticipated. However, the hi(J> prepane production due to high initial catalyst activity was not simulated. The actual unit heat balance was therefore not as predicted with the result that the process fired heater duty was significantly higher than expected. The startup heater which was in service to provide additional methanol superheat upstream of the DME reactor during startup was not taken off-line towards the end of startup as predicted. 

A simulation model for the reaction-regeneration steps in the transformation of methanol into hydrocarbons has been proposed and used for predicting the behaviour of a laboratory fixed bed pseudoadiabatic reactor. Kinetic models for both the main reaction and deactivation have been used, which take into account the attenuating role of water on both the zero time kinetics and the deactivation by coke deposition. The kinetics of coke combustion and the relationship between activity and coke content have been used for the design of the regeneration. The activity-coke content relationship is different in the reaction and regeneration steps. 

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