Catalytic reactors thermal conductivity

Van Driesen and Stewart (V4) have reported temperature measurements for various locations in commercial gas-liquid fluidized reactors for the large-scale catalytic desulfurization and hydrocracking of heavy petroleum fractions (2500 barrels per day capacity). The hydrogenation was carried out in two stages the maximum and minimum temperatures measured were 774° and 778°F for the first stage and 768° and 770°F for the second. These results indicate that gas-liquid fluidized reactors are characterized by a high effective thermal conductivity. 

Catalytic activity for the selective oxidation of H2S was tested by a continuous flow reaction in a fixed-bed quartz tube reactor with 0.5 inch inside diameter. Gaseous H2S, O2, H2, CO, CO2 and N2 were used without further purification. Water vapor (H2O) was introduced by passing N2 through a saturator. Reaction test was conducted at a pressure of 101 kPa and in the temperature range of 150 to 300 °C on a 0.6 gram catalyst sample. Gas flow rates were controlled by a mass flow controller (Brooks, 5850 TR) and the gas compositions were analyzed by an on-line gas chromotograph equipped with a chromosil 310 coliunn and a thermal conductivity detector. 

The catalytic reforming of CH4 by CO2 was carried out in a conventional fixed bed reactor system. Flow rates of reactants were controlled by mass flow controllers [Bronkhorst HI-TEC Co.]. The reactor, with an inner diameter of 0.007 m, was heated in an electric furnace. The reaction temperatoe was controlled by a PID temperature controller and was monitored by a separated thermocouple placed in the catalyst bed. The effluent gases were analyzed by an online GC [Hewlett Packard Co., HP-6890 Series II] equipped with a thermal conductivity detector (TCD) and carbosphere column (0.0032 m O.D. and 2.5 m length, 80/100 meshes), and identified by a GC/MS [Hewlett Packard Co., 5890/5971] equipped with an HP-1 capillary column (0.0002 m O.D. and 50 m length). 

The equations describing the concentration and temperature within the catalyst particles and the reactor are usually non-linear coupled ordinary differential equations and have to be solved numerically. However, it is unusual for experimental data to be of sufficient precision and extent to justify the application of such sophisticated reactor models. Uncertainties in the knowledge of effective thermal conductivities and heat transfer between gas and solid make the calculation of temperature distribution in the catalyst bed susceptible to inaccuracies, particularly in view of the pronounced effect of temperature on reaction rate. A useful approach to the preliminary design of a non-isothermal fixed bed catalytic reactor is to assume that all the resistance to heat transfer is in a thin layer of gas near the tube wall. This is a fair approximation because radial temperature profiles in packed beds are parabolic with most of the resistance to heat transfer near the tube wall. With this assumption, a one-dimensional model, which becomes quite accurate for small diameter tubes, is satisfactory for the preliminary design of reactors. Provided the ratio of the catlayst particle radius to tube length is small, dispersion of mass in the longitudinal direction may also be neglected. Finally, if heat transfer between solid cmd gas phases is accounted for implicitly by the catalyst effectiveness factor, the mass and heat conservation equations for the reactor reduce to [eqn. (62)]... 

The catalytic CO oxidation by pure oxygen was selected as a model reaction. The Pt/alumina catalyst In the form of 3.4 mm spherical pellets was used. The CO used In this study was obtained by a thermal decomposition of formic acid In a hot sulphuric acid. The reactor was constructed by three coaxial glass tubes. Through the outer jacket silicon oil was pumped, while air was blown through the inner jacket as a cooling medium. The catalyst was placed in the central part of the tube. The axial temperature profiles were measured by a thermocouple moving axially in a thermowell. Gas analysis was performed by an infrared analyzer or by a thermal conductivity cell. [7]. 

Let us now consider a catalytic packed bed reactor , i.e. a tubular reactor filled with a grained catalyst through which the gas mixture flows. With the particle diameter of the catalyst, dp, an additional dimensionless number dp/d is added to the pi-space the Reynolds number is now expediently formed with dp. The reaction rate is related to the unit of the bulk volume and characterized by an effective reaction rate constant ko,eff = k . The thermal conductivity (k) also has to be valid for the gas/bulk solids system and diffusion can be considered as being negligible (Sc is irrelevant). The complete pi-space is therefore ... 

As the model is to be constructed so that the intensive properties of the reacting fluid are to be invariant to the change of scale, such quantities as the heat capacity and the rate and heat of reaction are also invariant. In his treatment of packed catalytic reactors, Bosworth [see (B8), p. 318] assumes that the diffusivities and the thermal conductivity remain constant when the scale is changed. Since these quantities are approximately proportional to the mass velocity and the particle diameter, the resulting rules for scaling can not be correct. The presence of... 

By comparison with a fixed-bed gas-liquid reaction, a three-phase fluidized-bed reactor offers the advantage of very high effective thermal conductivity and, therefore, a more uniform temperature distribution in the reactor. Van Driesen and Stewart139 have demonstrated this for large-scale catalytic desulfurization and hydrocracking of heavy petroleum fractions. 

The low thermal conductivity of the molten polymers and their extremely high viscosity are the major problems for the catalytic cracking reactor design. The most widely used reactor systems have been ... 

While the above criteria are useful for diagnosing the effects of transport limitations on reaction rates of heterogeneous catalytic reactions, they require knowledge of many physical characteristics of the reacting system. Experimental properties like effective diffusivity in catalyst pores, heat and mass transfer coefficients at the fluid-particle interface, and the thermal conductivity of the catalyst are needed to utilize Equations (6.5.1) through (6.5.5). However, it is difficult to obtain accurate values of those critical parameters. For example, the diffusional characteristics of a catalyst may vary throughout a pellet because of the compression procedures used to form the final catalyst pellets. The accuracy of the heat transfer coefficient obtained from known correlations is also questionable because of the low flow rates and small particle sizes typically used in laboratory packed bed reactors. 

The catalytic reaction was performed in a fixed bed flow glass reactor. A gas mixture of CO2 and H2 (1 4, volume ratio) was passed continuously on the catalyst with F/W = 5400 ml g- h-i, unless otherwise mentioned. After the reaction the gas mixture was analyzed using a Chrompac MicroGC CP2002 gas chromatograph equipped with a thermal conductivity detector. 

The catalytic hydrogenation of carbon dioxide was performed in a continuous fixed bed reactor. The catalyst was reduced in a flow of hydrogen at 723 K for 20 - 24 hr. After the reduction, the catalyst was brought to the following conditions 573 K, 10 atm, space velocity of 1900 h-i and H2/CO2 = 3. The activity data was taken after 24h of reaction. The products were analyzed by a gas chromatograph (Chrompack CP 9001) equipped with thermal conductivity and flame ionization detectors. Carbon monoxide, carbon dioxide and water were analyzed on a Porapak Q column and the hydrocarbons on a GS Q capillary column. 

Catalytic activity was determined in a tubular packed b isothermal reactor at 500 K and 1 atm. A gas mixture was fed to the reactor at 350 cm min (CO 3%, HjO 26% Hj 48% N2 23% v/v) the catalyst weight was 0.04 g with a particle size of 0.177-0.250 mm. Reactants were analyzed by gas chromatography, using a thermal conductivity detector. Two packed columns were employed to analyze the reaction mixture. One was packed with 5A molecular sieve to separate hydrogen, nitrogen and CO, while COj was analyzed in a column packed with Porapak Q. Absence of diffusional control was experimentally verified by measuring the reaction conversion with catalyst particles of various sizes. The b was diluted (D%=10% v/v) with inert particles to provide isothermal conditions. 

Catalytic oxidation reactions were carried out in a conventional fwed bed reactor under atmospheric pressure [10]. The flow rate through the reactor was set at 500 cm min and the gas hourly space velocity (GHSV) was set at 15000 h. The residence time based on the packing volume of the catalyst was 0.24 s. Following the reactor, a portion of the effluent stream was delivered and analysed on-line using a Hewlett Packard 5890 Series II gas chromatograph (GC) equipped with an electron capture detector (ECD) and a thermal conductivity detector (TCD), and controlled with HP ChemStation software. The concentration of the chlorinated feeds was determined by the ECD after being separated in a HP-VOC column. 

The measurements of catalytic activity were carried out in a fixed bed reactor having 0.18 g of catalyst, in the temperature range 200-650-C and at a GHSV of 20000 h . The feed mixture to the reactor consisted of 1000 ppm of methane, 1000 ppm of NO, 2% of O2 and He as balance. The feed and effluent were analyzed by an on-line gas chromatograph using a column CTRl at 40°C and a thermal conductivity detector. 

The introduction of the catalyst presents one of the main problems in using MSRs for heterogeneously catalyzed reactions. There are some examples of reactors that are constructed directly from the catalytically active material. Kestenbaum et al. [145] used silver foils for the construction of a microchannel reactor for the partial oxidation of ethene to oxirane. A similar concept was proposed by Fichtner et al. [91,146], These authors used a microstructured rhodium catalyst for the partial oxidation of methane to syngas. This reaction can be considered as a coupling of the exothermic oxidation and the endothermic reforming of methane, which occur at different reaction rates. In such a case, the formation of a pronounced axial temperature profile can be avoided through the use of a material with high thermal conductivity. The reactor... 

Aligned multiwall CNT arrays were synthesized as a basis for a microstructured catalyst, which was then tested in the Fischer-Tropsch reaction in a microchannel reactor [269]. Fabrication of such a structured catalyst first involved MOCVD of a thin but dense A1203 film on a FeCrAlY foam to enhance the adhesion between the catalyst and the metal substrate. Then, multiwall CNTs were deposited uniformly on the substrate by controlled catalytic decomposition of ethene. Coating the outer surfaces of the nanotube bundles with an active catalyst layer results in a unique hierarchical structure with small interstitial spaces between the carbon bundles. The microstructured catalyst was characterized by the excellent thermal conductivity inherent to CNTs, and heat could be efficiently removed from the catalytically active sites during the exothermic Fischer-Tropsch synthesis. 

In contrast to catalytic treatment, coke formation during thermal treatment leads only to pipe blocking and poor thermal conductivity of the reactor walls. The most important difference between coke formation in catalytic and thermal treatment is that a free radical mechanism in coke formation is not possible in the case of thermal processing. 

The intense research effort carried out into the study of catalyst properties for the conversion of plastic wastes is in contrast with the few studies that have addressed reactor design. Thus, most of the studies use batch or simple fixed bed reactors despite the heat transfer and flow problems associated with the low thermal conductivity and high viscosity of the molten plastics. Various alternatives have been proposed to solve these problems the use of fluidized bed reactors, dissolution of the plastics in heavy oil fractions previously fed into the reactor, and a combination of thermal and catalytic treatments. However, all these processes present a number of difficulties, which makes further work on the reactor design necessary. 

It is obvious that catalytic distillation requires a reactor dedicated to one specific type of catalytic reaction. It can be questioned whether the fine-chemical industry performs many reactions needing a dedicated reactor. Application of the catalyst as a thin porous layer on the surface of a metal or ceramic material, however, affords interesting possibilities in the fine-chemical industry. With liquid-phase reactions, the catalyst is still almost completely involved in the reaction when the layer in which the catalyst is present is not much thicker than ca 100 pm. Separation of the catalyst from the reaction product, removal of the catalyst from the reactor, and collection and storage of the catalyst is no longer required this greatly facilitates operation. The catalyst can, furthermore, be treated thermally in a gas flow, because the pressure drop depends on the structure of the solid on which the catalyst has been applied. This structure can easily be selected thus that the pressure drop is low. When, finally, the catalyst is applied to a metal surface with appreciable thermal conductivity the temperature of the reaction can be maintained accurately at the value desired. 

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