Reactor selection, benzene alkylation

Thus, in ammonia synthesis, mixed oxide base catalysts allowed new progress towards operating conditions (lower pressure) approaching optimal thermodynamic conditions. Catalytic systems of the same type, with high weight productivity, achieved a decrease of up to 35 per cent in the size of the reactor for the synthesis of acrylonitrile by ammoxidation. Also worth mentioning is the vast development enjoyed as catalysis by artificial zeolites (molecular sieves). Their use as a precious metal support, or as a substitute for conventional silico-aluminaies. led to catalytic systems with much higher activity and selectivity in aromatic hydrocarbon conversion processes (xylene isomerization, toluene dismutation), in benzene alkylation, and even in the oxychlorination of ethane to vinyl chloride. 

Three operation variables, benzene-to-propylene ratio (Rg), temperature (T), and space time (t) are of prime concern to a reactor designer. Since alkylation reactions are so fast, the controlling performance of the reactor is determined by the reaction selectivity. At a constant average temperature, examination of the reaction paths and kinetics models indicated that selectivity Increased with Increasing benzene-to-propylene ratio regardless of the space time applied. The temperature effect was slightly more complicated. Since most of a process reactor... 

Design a reactor for the alkylation of benzene with propylene to maximize the selectivity of isopropylbenzene. [Prac. 2nd Joint China/USA Chem. Eng. Conf III, 51, (1997)]. 

The benzene—toluene fraction is further fractionated in a small column, not shown in Figure 5, to recover benzene for recycle to the alkylation unit and toluene for sale. This toluene can be converted to benzene by hydrodealkylation but the high selectivity catalyst has reduced the formation of toluene in the dehydrogenation reactor to the point where the cost of installing a hydrodealkylation unit is difficult to justify even in a large styrene plant. 

The LAB production process (process 1) is mainly developed and licensed by UOP. The N-paraffins are partially converted to internal /z-olefins by a catalytic dehydrogenation. The resulting mixture of /z-paraffins and n-olefins is selectively hydrogenated to reduce diolefins and then fed into an alkylation reactor, together with an excess benzene and with concentrated hydrofluoric acid (HF) which acts as the catalyst in a Friedel-Crafts reaction. In successive sections of the plant the HF, benzene, and unconverted /z-paraffins are recovered and recycled to the previous reaction stages. In the final stage of distillation, the LAB is separated from the heavy alkylates. 

Catalyst deactivation due to coke formation is relatively speedy for a reactions such as alkylation of benzene to ethylbenzene over zeolite, particularly when the benzene to ethylene ratio is low. Another problem of this reaction is the formation of xylenes, the major byproducts. Though their total amount produced in the process is very limited, they are harmful to the process because of the difficulty to remove them from the desired product ethylbenzene. Therefore, investigating the mechanism of catalyst coking is of practical significance for finding the potential ways for prolonging the reactor runtime and decreasing the xylenes selectivity. 

The monoalkylation selectivity of the alkylation step refers to the fraction of ethylene that reacts to form ethylbenzene, as opposed to forming polyethylated species. To suppress the formation of PEBs, benzene must be fed to the alkylation reactor in considerable excess (frequently five to seven times the stoichiometric requirement). Equipment in the alkylation reaction and benzene recovery systems must therefore be sized to accommodate the flow of excess benzene, and energy must be expended to recover the excess benzene from the reactor effluent. However, the superior monoalkylation selectivity and stability of MCM-22 permits operation with reduced benzene feed rates - in the range of two to four times the stoichiometric requirement - without excessive PEB formation (see Table 11.1). 

Description Linear paraffins are fed to a Pacol reactor (1) to dehydrogenate the feed into corresponding linear olefins. Reactor effluent is separated into gas and liquid phases in a separator (2). Diolefins in the separator liquid are selectively converted to mono-olefins in a DeFine reactor (3). Light ends are removed in a stripper (4) and the resulting olefin-paraffin mixture is sent to a Detal reactor (5) where the olefins are alkylated with benzene. The reactor effluent is sent to a fractionation section (6, 7) for separation and recycle of unreacted benzene to the Detal reactor, and separation and recycle of unreacted paraffins to the Pacol reactor. A rerun column (8) separates the LAB product from the heavy alkylate bottoms stream. 

Figure 4.10 Typical results of the calculations of expected variations in the stationary concentrations of components at the benzene (B) alkylation with ethylene (E) along a plug-flow reactor of length L at 210 C x is the distance from the inlet of the reactor. The calculations were performed in terms of the Horiuti-Boreskov-Onsager reciprocity relations to optimize the composition of the initial reaction mixture so the outlet and inlet diethylbenzene (DBE) concentrations would be identical, which means 100% selectivity of the process in respect to the target product ethylbenzene (EB).

The UOP Paeol process for selective long-chain paraffin dehydrogenation to produee linear mono-olefins is shown in Fig. 15 in combination with the UOP detergent alkylation process. The Pacol process consists of a radial-flow reactor and a product recovery section. Worldwide, more than 2 million metric tons per year of linear alkyl benzene is produced employing this process. 

The oxidation of various hydrocarbons such as n-octane, cyclohexane, toluene, xylenes and trimethyl benzenes over two vanadium silicate molecular sieves, one a medium pore VS-2 and the other, a novei, iarge pore V-NCL-1, in presence of aqueous HjOj has been studied. These reactions were carried out in batch reactors at 358-373 K using acetonitrile as the solvent. The activation of the primary carbon atoms in addition to the preferred secondary ones in n-octane oxidation and oxidation of the methyl substituents in addition to aromatic hydroxyiation of alkyl aromatics distinguish vanadium silicates from titanium silicates. The vanadium silicates are also very active in the secondary oxidation of alcohols to the respective carbonyl compounds. V-NCL-1 is active in the oxidation of bulkier hydrocarbons wherein the medium pore VS-2 shows negligible activity. Thus, vanadium silicate molecular sieves offer the advantage of catalysing selective oxidation reactions in a shape selective manner. 

Modem alkylation processes make use of solid catalysts based on zeolites. According to different technologies, the reaction can be performed in vapour or liquid phase. The selection of a suitable chemical reactor for ethylbenzene is discussed in the Example 8.3. A conceptual flowsheet is depicted in Fig. 7.31 for a vapour-phase process (Mobil-Badger), one of the most widely used. The reactor works at 390-440 C and 0.6-3 MPa. Besides the main product ethylbenzene (EB), polyethylbenzenes (PEB) are formed, their amount depending on the reaction conditions. Large excess of benzene, over 6 1, is needed to shift the equilibrium to the desired product. The reaction mixture is sent to the separation section. Final yield can increase over 99% by converting PEB s to EB in a separate transalkylation reactor. 

Ethylbenzene (EB) is currently produced by alkylation of benzene with ethylene, primarily via two routes liquid-phase with AlCl, catalyst, or vapour-phase in catalytic fixed bed reactor (Ullmann, 2001). Examine the differences, as well as advantages and disadvantages of these routes. List pros and cons in selecting suitable reactors. 

The gas phase alkylation of toluene with methanol was carried out in a fixed-bed tubular reactor at atmospheric pressure. Samples were sieved to retain particles with 0.35-0.40 mm in diameter for catafytic measurements. A mixture of toluene/methanol of 1 1 molar ratio was vaporized in a preheating section and delivered to the reactor. The reaction was carried out at 400 °C, employing a space velocity (WHSV) of 2 h. Toluene conversion (Xtoi) was calculated as Xtoi (%) = [EYj / (ZYj + Ytoi.)]100, where ZYj is the molar fiactions of the aromatic reaction products, including benzene, and Ytoi is the outlet molar fiaction of toluene. The selectivity to product j was determined as Sj (%) = [Y/EYj.lOO. The Sei-bz selectivity includes the sum of ethylbenzene and styrene, which are the side-chain alkylation products. In-situ poisoning experiments were carried out by doping the toluene/methanol mixture with either acetic acid or 3,5-dimethyl pyridine in a concentration range between 0-15000 ppm... 

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