Catalytic catalyst beds

Catalytic methanation processes include (/) fixed or fluidized catalyst-bed reactors where temperature rise is controlled by heat exchange or by direct cooling using product gas recycle (2) through wall-cooled reactor where temperature is controlled by heat removal through the walls of catalyst-filled tubes (J) tube-wall reactors where a nickel—aluminum alloy is flame-sprayed and treated to form a Raney-nickel catalyst bonded to the reactor tube heat-exchange surface and (4) slurry or Hquid-phase (oil) methanation. 

The Claus process is the most widely used to convert hydrogen sulfide to sulfur. The process, developed by C. F. Claus in 1883, was significantly modified in the late 1930s by I. G. Farbenindustrie AG, but did not become widely used until the 1950s. Figure 5 illustrates the basic process scheme. A Claus sulfur recovery unit consists of a combustion furnace, waste heat boiler, sulfur condenser, and a series of catalytic stages each of which employs reheat, catalyst bed, and sulfur condenser. Typically, two or three catalytic stages are employed. 

Process Description. Reactors used in the vapor-phase synthesis of thiophene and aLkylthiophenes are all multitubular, fixed-bed catalytic reactors operating at atmospheric pressure, or up to 10 kPa and with hot-air circulation on the shell, or salt bath heating, maintaining reaction temperatures in the range of 400—500°C. The feedstocks, in the appropriate molar ratio, are vaporized and passed through the catalyst bed. Condensation gives the cmde product mixture noncondensable vapors are vented to the incinerator. 

Cost. The catalytically active component(s) in many supported catalysts are expensive metals. By using a catalyst in which the active component is but a very small fraction of the weight of the total catalyst, lower costs can be achieved. As an example, hydrogenation of an aromatic nucleus requires the use of rhenium, rhodium, or mthenium. This can be accomplished with as fittie as 0.5 wt % of the metal finely dispersed on alumina or activated carbon. Furthermore, it is almost always easier to recover the metal from a spent supported catalyst bed than to attempt to separate a finely divided metal from a liquid product stream. If recovery is efficient, the actual cost of the catalyst is the time value of the cost of the metal less processing expenses, assuming a nondeclining market value for the metal. Precious metals used in catalytic processes are often leased. 

Design nd Operation. The destruction efficiency of a catalytic oxidation system is determined by the system design. It is impossible to predict a priori the temperature and residence time needed to obtain a given level of conversion of a mixture in a catalytic oxidation system. Control efficiency is determined by process characteristics such as concentration of VOCs emitted, flow rate, process fluctuations that may occur in flow rate, temperature, concentrations of other materials in the process stream, and the governing permit regulation, such as the mass-emission limit. Design and operational characteristics that can affect the destmction efficiency include inlet temperature to the catalyst bed, volume of catalyst, and quantity and type of noble metal or metal oxide used. 

The size of the catalyst bed depends mainly on the degree of VOC reduction requited (14). VOC destmction efficiencies up to 95% can usually be attained using reasonable space velocities (14). However, the low GHSVs, and subsequently high catalyst volumes requited to achieve extremely high (eg, 99%) conversions, can sometimes make catalytic oxidation uneconomical. Conventional bed geometries may be found in the Hterature (14). 

The various reaction rate properties of the different solvents influence the design of a catalytic reactor. Eor example, for a specific catalyst bed design, an effluent stream containing a preponderance of monohydric alcohols, aromatic hydrocarbons, or propjiene requires a lower catalyst operating temperature than that required for solvents such as isophorone and short-chain acetates. 

The catalytic decomposition of acetylene was carried out in a flow reactor at atmospheric pressure. A ceramic boat containing 20-100 mg of the catalyst was placed in a quartz lube (inner diameter 4-10 mm, length 60-100 cm). The reaction mixture of 2.5-10% C2H2 (Alphagaz, 99.6%) in Nj (Alphagaz, 99.99%) was passed over the catalyst bed at a rate of 0.15-0.59 mol C2H2 g h for several hours at temperatures in the range 773-1073 K. 

In catalytic incineration, there are limitations concerning the effluent streams to be treated. Waste gases with organic compound contents higher than 20% of LET (lower explosion limit) are not suitable, as the heat content released in the oxidation process increases the catalyst bed temperature above 650 °C. This is normally the maximum permissible temperature to which a catalyst bed can be continuously exposed. The problem is solved by dilution-, this method increases the furnace volume and hence the investment and operation costs. Concentrations between 2% and 20% of LET are optimal, The catalytic incinerator is not recommended without prefiltration for waste gases containing particulate matter or liquids which cannot be vaporized. The waste gas must not contain catalyst poisons, such as phosphorus, arsenic, antimony, lead, zinc, mercury, tin, sulfur, or iron oxide.(see Table 1.3.111... 

Catalytic combustor A device used to remove various solid, liquid, or gaseous pollutants from air or another gas, in which the gas is heated by an open burner to between 250 and 500 °C and passed through a catalyst bed in which the organic contaminants are oxidized into harmless by-products. 

Catalytic reformers are normally designed to have a series of catalyst beds (typically three beds). The first bed usually contains less catalyst than the other beds. This arrangement is important because the dehydrogenation of naphthenes to aromatics can reach equilibrium faster than the other reforming reactions. Dehydrocyclization is a slower reaction and may only reach equilibrium at the exit of the third reactor. Isomerization and hydrocracking reactions are slow. They have low equilibrium constants and may not reach equilibrium before exiting the reactor. 

Zabor et al. (Zl) have described studies of the catalytic hydration of propylene under such conditions (temperature 279°C, pressure 3675 psig) that both liquid and vapor phases are present in the packed catalyst bed. Conversions are reported for cocurrent upflow and cocurrent downflow, it being assumed in that paper that the former mode corresponds to bubble flow and the latter to trickle-flow conditions. Trickle flow resulted in the higher conversions, and conversion was influenced by changes in bed height (for unchanged space velocity), in contrast to the case for bubble-flow operation. The differences are assumed to be effects of mass transfer or liquid distribution. 

In the case of a catalytic membrane reactor (CMR), the membrane is (made) intrinsically catalytically active. This can be done by using the intrinsic catalytic properties of the zeolite or by making the membrane catalytically active. When an active phase is deposited on top of a membrane layer, this is also called a CMR because this becomes part of the composite membrane. In addition to the catalytic activity of the membrane, a catalyst bed can be present (PBCMR). The advantages of a CMR are as follows ... 

When NOj levels are measured electrochemicaUy, NO and NO2 can lead to opposing signals because NO is oxidized and NO2 tends to be reduced. Moreover, it is preferred to obtain a total NO, measurement instead of only one of the constituents. The latter can be achieved by catalytically equilibrating the feed with oxygen before contact with the sensor by coating an active zeolite layer on top or placing a active catalyst bed in front of the sensor. Both approaches have been demonstrated successfully with a Pt-Y zeohte as active catalyst [74, 75]. The additional advantage of the filter bed is a reduction in the cross-sensitivity with CO due to CO oxidation above 673 K. 

Traditionally, an average Sherwood number has been determined for different catalytic fixed-bed reactors assuming constant concentration or constant flux on the catalyst surface. In reality, the boundary condition on the surface has neither a constant concentration nor a constant flux. In addition, the Sh-number will vary locally around the catalyst particles and in time since mass transfer depends on both flow and concentration boundary layers. When external mass transfer becomes important at a high reaction rate, the concentration on the particle surface varies and affects both the reaction rate and selectivity, and consequently, the traditional models fail to predict this outcome. 

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 most important undesired metallic impurities are nickel and vanadium, present in porphyrinic structures that originate from plants and are predominantly found in the heavy residues. In addition, iron may be present due to corrosion in storage tanks. These metals deposit on catalysts and give rise to enhanced carbon deposition (nickel in particular). Vanadium has a deleterious effect on the lattice structure of zeolites used in fluid catalytic cracking. A host of other elements may also be present. Hydrodemetallization is strictly speaking not a catalytic process, because the metallic elements remain in the form of sulfides on the catalyst. Decomposition of the porphyrinic structures is a relatively rapid reaction and as a result it occurs mainly in the front end of the catalyst bed, and at the outside of the catalyst particles. 

In practice, granular beds comprising a very large number of catalyst pellets are used. It is well known that the efficiency of a catalytic reactor depends crucially on the liquid phase distribution within the catalyst bed [14]. It is likely that the development of hot spots in a catalyst bed is also related to the character of liquid phase distribution. Therefore, it is very important to map the spatial distribution of the liquid phase in a catalytic reactor for various operation regimes. This eventually should lead to the formulation of the mechanisms responsible for the development of critical phenomena on both a micro- and macroscale. 

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