Fixed-bed reactors operation

The thermal catalytic route proposed involves heating the fresh reactant feed plus recycle up to 790°C and feeding this material into a M0S2 catalyst fixed-bed reactor operating at 0.1 MPa (1 atm). The route yields a production of H2 almost 50% higher than the decomposition of H2S route. 

The Catofin process, which was formerly the property of Air Products (Houdry Division), uses a proprietary chromium catalyst in a fixed-bed reactor operating under vacuum. There are actually multiple reactors operating in cycHc fashion. In sequence, these reactors process feed for about nine minutes and are then regenerated for nine minutes. The chromium catalyst is reduced from Cr to Cr during the regeneration cycle. 

In principle, the catalytic converter is a fixed-bed reactor operating at 500—620°C to which is fed 200—3500 Hters per minute of auto engine exhaust containing relatively low concentrations of hydrocarbons, carbon monoxide, and nitrogen oxides that must be reduced significantly. Because the auto emission catalyst must operate in an environment with profound diffusion or mass-transfer limitations (51), it is apparent that only a small fraction of the catalyst s surface area can be used and that a system with the highest possible surface area is required. 

Despite the higher cost compared with ordinary catalysts, such as sulfuric or hydrochloric acid, the cation exchangers present several features that make their use economical. The abiHty to use these agents in a fixed-bed reactor operation makes them attractive for a continuous process (50,51). Cation-exchange catalysts can be used also in continuous stirred tank reactor (CSTR) operation. 

The Powerformer reaction absorbs heat it is largely an endothermic reaction. This heat of reaction is in the order of 200-350 BTU/pound, depending on the type of feed. Since the individual Powerforming fixed bed reactors operate... 

Few fixed-bed reactors operate in a region where the intrinsic kinetics are applicable. The particles are usually large to minimize pressure drop, and this means that diffusion within the pores. Steps 3 and 7, can limit the reaction rate. Also, the superficial fluid velocity may be low enough that the external film resistances of Steps 2 and 8 become important. A method is needed to estimate actual reaction rates given the intrinsic kinetics and operating conditions within the reactor. The usual approach is to define the effectiveness factor as... 

A packed-bed nonpermselective membrane reactor (PBNMR) is presented by Diakov et al. [31], who increased the operational stability in the partial oxidation of methanol by feeding oxygen directly and methanol through a macroporous stainless steel membrane to the PB. Al-Juaied et al. [32] used an inert membrane to distribute either oxygen or ethylene in the selective ethylene oxidation. By accounting for the proper kinetics of the reaction, the selectivity and yield of ethylene oxide could be enhanced over the fixed-bed reactor operation. 

Acetic acid is formed when methane reacts with CO or C02 in aqueous solution in the presence of 02 or H202 catalyzed by vanadium complexes.327 A Rh-based FeP04 catalyst applied in a fixed-bed reactor operating at atmospheric pressure at 300-400° C was effective in producing methyl acetate in the presence of nitrous oxide.328 The high dispersion of Rh at sites surrounded by iron sites was suggested to be a key factor for the carbonylation reaction. 

F. Larachi, A. Laurent, G. Wild and N. Midoux, Some experimental liquid saturation results in fixed-bed reactors operated under elevated pressure in cocurrent upflow and downflow of the gas and the liquid, Ind. Engng. Chem. Res., 30 (1991) 2404-2410. 

The o-xylene oxidation was carried out in a continuous flow fixed bed reactor operating at atmospheric pressure. The feed mixture (0.7 mol%) was obtained injecting the organic reactant in the air flow. 

Figure 1. (a) Step-wise function of inlet parameters change, (b) Fixed-bed reactor operated at periodic switching between two dilTerent feeds that provide for step-wise control. 

Figure 2. Schemes of fixed-bed reactors operated under forced unsteady-state conditions (a) Reverse-flow reactor (b) Rotary reactor (c) Reactor system with periodic changes between the inlet and outlet ports in two fixed beds. The tables show positions of switching valves during two successive cycles C = valve closed O = valve open.

Poisoning Effects in Temperature-Increased Fixed-Bed Reactor Operation... 

Description The ATOFINA/UOP Olefin Cracking Process was jointly developed by Total Petrochemicals (formerly ATOFINA) and UOP to convert low-value C4 to C8 olefins to propylene and ethylene. The process features fixed-bed reactors operating at temperatures between 500°C and 600°C and pressures between 1 and 5 bars gauge. 

Marin et al. (250) attempted to model a reactor similar to that used by Alonso and co workers. Their simulations were compared with simulations representing a fixed-bed reactor operated under similar conditions. They concluded that the membrane reactor (with the external fluidized bed) was a viable technology for n-butane oxidation, but that it offered only a modest increase in MA yields relative to those realized in a fixed-bed reactor. Nonetheless, the safer operating conditions which keep the O2 and hydrocarbon flows separate, particularly with the oxidation of butane to MA, are desirable. Presently, MA yields are chiefly governed by the explosive limits of butane in air (i.e., 1.8%). Increasing the butane concentration with an optimized membrane reactor may increase overall MA yields. 

Figure 1. Schematic of radial flow, fixed bed reactor operation...

Catalytic dealkylation of cumene was carried out in a fixed bed reactor operated at 350 C and atmospheric pressure. The contact time was 1.5 sec and the WHSV was 0.4g cumene/g/hr. The conversion of cumene was determined using a Hewlett Packard 5890 gas Chromatograph equipped with a Supelco wide bore capillary column. 

An analysis is made of the factors which pose a limit to representative downscaling of catalyst testing in continuous fixed-bed reactors operated with either gas or gas-liquid flow. Main limiting factors are the axial dispersion and, in the case of gas-liquid operation, also the contacting of the catalyst. The effects of catalyst and reactor geometries are quantified, and boundaries for safe operation are indicated. 

Catalytic measurements. The catalytic tests were performed in fixed bed reactors operating at 463-498 K and total pressure of 1-20 bar. The H2/CO ratio was 2 in all experiments. Prior to the reaction, the catalysts were reduced in the flow of hydrogen at 753-773 K for 5 h. FT catalytic rates and selectivities were measured at the stationary regime after 24 h time-on-stream. FT reaction rates were normalized by the number of cobalt auims in the reactor. The reaction products were analyzed by gas chromatography. 

Realistic analysis of fixed-bed reactor experiments requires calculation of interfacial states. Laboratory reactors are typically much shorter than full-scale units and operate with smaller axial velocities, producing significant departures of the iiiterfacial states from the measurable values in the mainstream fluid and consequent difficulties in establishing catalytic reaction models. Interfacial temperatures and partial pressures were calculated with jn and jo- and used in estimating reaction model parameters, in a landmark paper by Yoshida. Ramaswami. and Hougen (1962). Here we give an updated analysis of interfacial states in fixed-bed reactor operations for improved treatment of catalytic reaction data. 

The isomerization experiments were carried out in a fixed bed reactor operating at 3.0 MPa total pressure and using a simulated LSR feed formed by mixture of n-pentane/n-hexane (60/40 by weight). The reaction conditions were temperature (523-583 K), hydrogen/feed ratio (1.5-11 mol/mol) and WHSV (1.8-8.0 h ). Before the experiments Pt/beta catalysts were reduced in situ with a flow of 350 cm of hydrogen at 723 K for two hours. Product analyses were performed on line in a gas chromatrograph equipped with a 25m x 0.32mm ID fused silica column coated with poraplot U and a FID. 

Frycek, G, J. Experiment and Modeling of Poisoning Effects in Temperature-Increased Fixed Bed Reactor Operation, Ph D Dissertation, Northwestern University, Evanston, IL, 1984. Available from University Hicrofilms, Inc. 

The catalyst, 50% M/H-ZSM-5 and 50% alumina, was supplied by INTEVEP S.A. as 1/16 inch diameter cylinders. Its catalytic properties (activity, yield and selectivity) was studied in an isothermal fixed bed reactor operated under hydrogen atmosphere keeping the ratio of H/n-C8 always equal to 42.06. Deactivation tests were done under hydrogen atmosphere and at the following reaction conditions WWH = 2.87 h, T = 280°C, P = 200 psig., until a stable activity level or a complete deactivation was reached. For pure n-octane transformation over a stabilized catalyst, the change in products distribution with contact time, fi om 0.2 to 2 h, and with reaction temperatures, fi om 200 to 470°C, was also studied [9]. 

To optimize the performance of the fixed bed reactor operation several constructions of fixed bed reactors have been investigated over the years. Three of the most common reactor designs are ... 

Periodic flow reversal inducing forced unsteady-state conditions [339]. The flow to the reactor is continuously reversed before the steady state is attained. A dual hot-spot temperature profile, characterized by a considerably lower temperature than in the single hot spot that would develop in the traditional flow configuration, forms in exothermic oxidation reactions. An increase in selectivity and better reactor control (lower risk of runaway) is possible over fixed-bed reactor operations, but compared... 

The oxidation is carried out in fixed bed reactors, operating at 6 atm under adiabatic conditions with an inlet temperature of 95 °C and outlet temperature of 110 °C. The overall oxidation section performances per pass are ... 

After separation of MIBK by distillation, the by-products, obtained in aqueous solution, are fed to the HDO section for selective hydrogenation. The HDO reaction is carried out in fixed bed reactors, operating under adiabatic conditions at 400 °C inlet temperature and 25 bar of hydrogen pressure. Typically, the hydrogen/dihydroxy benzenes molar ratio is set to 20. The produced phenol is recovered by distillation and recycled to the process cycle, thus avoiding any coproduction of dihydroxybenzenes. 

When a fixed-bed reactor operates at steady state, an amount of energy equal to the heat released by reaction on the catalyst pellet must be transferred to the bulk fluid. In Sec. 10-2 this requirement was used to relate the concentration and temperature differences between pellet and fluid. Here we want to develop a method for predicting the magnitude of the temperature... 

Activity measurements were performed in a continuous flow fixed bed reactor operating at atmospheric pressure between 100 °C and 400 °C. The total flow rate was 100 cc/min and feed composition was NO (0,1 vol. %), NH3 (0,11 vol. %) and O2 (2,5 vol. %), in helium. The catalyst amount of 0,150 g calcined at 400 °C and the space velocity were kept constant for all experiments. The inlet and outlet gas compositions were measured using a quadrupole mass spectrometer QMC 311 Balzers coupled to the reactor. 

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