Plants fixed bed reactors

A dynamic experimental method for the investigation of the behaviour of a nonisothermal-nonadiabatic fixed bed reactor is presented. The method is based on the analysis of the axial and radial temperature and concentration profiles measured under the influence of forced uncorrelated sinusoidal changes of the process variables. A two-dimensional reactor model is employed for the description of the reactor behaviour. The model parameters are estimated by statistical analysis of the measured profiles. The efficiency of the dynamic method is shown for the investigation of a pilot plant fixed bed reactor using the hydrogenation of toluene with a commercial nickel catalyst as a test reaction. 

In the present work a method is described to extract the information necessary for modelling from only a few dynamic experimental runs. The method is based on the measurement of the changes of the temperature and concentration profiles in the reactor under the influence of forced simultaneous sinusoidal variations of the process variables. The characteristic features of the dynamic method are demonstrated using the behaviour of a nonisothermal-nonadiabatic pilot plant fixed bed reactor as an example. The test reaction applied was the hydrogenation of toluene to methylcyclohexane on a commercial nickel catalyst. 

A. Baiker, D. Epple and A. Wokaun, Behaviour of a pilot plant fixed bed reactor during catalyst deactivation, Chem.Eng,Sci.,41(1986)779. 

Zein el Deen et al. (30) studied the kinetics of the FTS on sintered oxides of iron and manganese. They observed, too, that the rate is independent of the CO partial pressure. Bub et al. (20) developed empirical expressions for the production rate of CO2 and Ci to C4 hydrocarbons on a Mn/Fe catalyst which could be used to successfully describe the conversion and selectivity in a pilot plant fixed bed reactor (2 cm ID by 80 cm length). If a catalyst like Mn/Fe gives a Schulz-Flory product distribution the hydrocarbon fraction can be calculated from the overall conversion rate and the chain growing probability a... 

Pilot plant fixed-bed reactors are traditionally designed at space velocities equivalent to commercial scale fixed-bed reactors. Thus the film diffusion resistance of the process at the two scales is different. In general, pilot plant fixed-bed reactors are film diffusion rate limited, whereas commercial-size fixed-bed reactors are either pore diffusion rate or, more rarely, reaction rate limited. This shift from film diffusion rate limited to, more generally, pore diffusion rate limited occurs due to the high volumetric fluid flow through the catalyst mass in a commercial-size fixed-bed reactor. Thus reactant consumption or product formation is faster in the commercial-size fixed-bed reactor than in the pilot plant fixed-bed reactor. 

Data for the production and sales of maleic anhydride and fumaric acid ia the United States between 1979 and 1992 are shown ia Table 5. Production of maleic anhydride during this time grew - 2% on average per year. Production of fumaric acid has declined during the same period as customers have switched to the less cosdy maleic anhydride when possible. All production of maleic anhydride in the United States in 1992 was from butane-based plants which used fixed-bed reactor technology as shown in Table 6. The number of fumaric acid producers has been reduced considerably since the early 1980s with only two producers left in the United States in 1992 as shown in Table 6. Pfizer shut down its fumaric acid plant at the end of 1993. However, Bartek of Canada will start up an expanded fumaric acid faciUty to supply the North American market for both their own and Huntsman s requirements. 

The Fischer-Tropsch reaction is highly exothermic. Therefore, adequate heat removal is critical. High temperatures residt in high yields of methane, as well as coking and sintering of the catalyst. Three types of reac tors (tubular fixed bed, fluidized bed, and slurry) provide good temperature control, and all three types are being used for synthesis gas conversion. The first plants used tubular or plate-type fixed-bed reactors. Later, SASOL, in South Africa, used fluidized-bed reactors, and most recently, slurry reactors have come into use. 

Shell Gas B.V. has constructed a 1987 mVd (12,500 bbhd) Fischer-Tropsch plant in Malaysia, start-up occurring in 1994. The Shell Middle Distillate Synthesis (SMDS) process, as it is called, uses natural gas as the feedstock to fixed-bed reactors containing cobalt-based cat- yst. The heavy hydrocarbons from the Fischer-Tropsch reactors are converted to distillate fuels by hydrocracking and hydroisomerization. The quality of the products is very high, the diesel fuel having a cetane number in excess of 75. 

So far, consideration has been limited to chemistry physical constraints such as heat transfer may also dictate the way in which reactions are performed. Oxidation reactions are highly exothermic and effectively there are only two types of reactor in which selective oxidation can be achieved on a practical scale multitubular fixed bed reactors with fused salt cooling on the outside of the tubes and fluid bed reactors. Each has its own characteristics and constraints. Multitubular reactors have an effective upper size limit and if a plant is required which is too large to allow the use of a single reactor, two reactors must be used in parallel. 

Georgiou D, Hatiras J, Aivasidis A (2005) Microbial immobilization in a two stage fixed bed reactor pilot plant for onsite anaerobic decolorization of textile wastewater. Enzyme Microb Technol 37 597-605... 

The asymptotic observer was tested in an experimental 1 m upflow anaerobic fixed bed reactor pilot plant used for the treatment of industrial wine distillery vinasses obtained from local distilleries in the Narbonne area (France) (see Figure 26). These experimental runs were carried out over a 35 day period. Measurements of the dilution rate as well as the and S2 concentrations and the partial CO2 pressure were performed on-line (see Figures 27 to 30). The measurements of S J, and were obtained from off-line data and... 

The interval observer was finally tested on-line in a series of experimental runs conducted over a 35 days period in the anaerobic fixed bed reactor pilot plant. Parameters used are listed in Table 4. For the experimental implementation of the observer the influent concentrations were considered unknown and only a known boundary region was supplied to the interval observer. Figures 50 to 52 depict the corresponding bounded intervals for S, S2 and Z , respectively. 

Fixed bed plants. In this type of plant, the process flow for all three feeds looks like the plant in Figure 20—3. The feed and compressed air are mixed, vaporized in a heater, and then charged to the fixed-bed reactor, a bundle of rubes packed with the catalyst. The ratio of air to hydrocarbon is generally about 75 1 to keep the mixture outside the explosive range, always a good idea. The feed temperature is 800-900°F, depending on the feed. The reaction time is extremely quick, so the feed is in contact with the catalyst for only 0.1 to 1.0 second. 

Another difference between Co and Fe is their sensitivity towards impurities in the gas feed, such as H2S. In this respect, Fe-based catalysts have been shown to be more sulfur-resistance than their Co-based counterparts. This is also the reason why for Co F-T catalysts it is recommended to use a sulphur-free gas feed. For this purpose, a zinc oxide bed is included prior to the fixed bed reactor in the Shell plant in Malaysia to guarantee effective sulphur removal. Co and Fe F-T catalysts also differ in their stability. For instance, Co-based F-T systems are known to be more resistant towards oxidation and more stable against deactivation by water, an important by-product of the FTS reaction (reaction (1)). Nevertheless, the oxidation of cobalt with the product water has been postulated to be a major cause for deactivation of supported cobalt catalysts. Although, the oxidation of bulk metallic cobalt is (under realistic F-T conditions) not feasible, small cobalt nanoparticles could be prone to such reoxidation processes. 

Ethyl Chloride. Hydrochlorination of ethylene with HC1 is carried out in either the vapor or the liquid phase, in the presence of a catalyst.182-184 Ethyl chloride or 1,2-dichloroethane containing less than 1% A1C13 is the reaction medium in the liquid-phase process operating under mild conditions (30-90°C, 3-5 atm). In new plants supported AlClj or ZnCl2 is used in the vapor phase. Equimolar amounts of the dry reagents are reacted in a fluidized- or fixed-bed reactor at elevated temperature and pressure (250-400°C, 5-15 atm). Both processes provide ethyl chloride with high (98-99%) selectivity. 

The effect of ZSM-5 as an octane additive to a cracking catalyst was studied in both a small fixed-bed reactor and a fluidized-bed pilot plant. Analyses of the products of these tests were used to determine the reaction chemistry. It was found that by maximizing the ratio of isomerization activity to hydrogen transfer activity, gasoline octane was increased with a minimum of yield loss. This could be accomplished by increasing the silica-to-alumina ratio of the additive zeolite. 

Of the indirect liquefaction procedures, methanol synthesis is the most straightforward and well developed [Eq. (6)]. Most methanol plants use natural gas (methane) as the feedstock and obtain the synthesis gas by the steam reforming of methane in a reaction that is the reverse of the methanation reaction in Eq. (5). However, the synthesis gas can also be obtained by coal gasification, and this has been and is practiced. In one modern low-pressure procedure developed by Imperial Chemical Industries (ICI), the synthesis gas is compressed to a pressure of from 5 to 10 MPa and, after heating, fed to the top of a fixed bed reactor containing a copper/zinc catalyst. The reactor temperature is maintained at 250 to 270°C by injecting... 

All of these applications use fixed-bed reactors, and, very importantly, all were scaled up from bench-scale pilot plant data. This successful scale-up experience with the ZSM-5 catalyst was an important consideration in formulating the MTG development strategy. 

In some cases a plant may have a pre-reformer. A pre-former is an adiabatic, fixed-bed reactor upstream of the primary reformer. It provides an operation with increased flexibility in the choice of feed stock it increases the life of the steam reforming catalyst and tubes it provides the option to increase the overall plant capacity and it allows the reformer to operate at lower steam-to-carbon ratios166. The hot flue gas from the reformer convection section provides the heat required for this endothermic reaction. 

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