Catalysts VOLUME

Adl b tic Converters. The adiabatic converter system employs heat exchangers rather than quench gas for interbed cooling (Fig. 7b). Because the beds are adiabatic, the temperature profile stiU exhibits the same sawtooth approach to the maximum reaction rate, but catalyst productivity is somewhat improved because all of the gas passes through the entire catalyst volume. Costs for vessels and exchangers are generally higher than for quench converter systems. 

Stea.m-Ra.ising Converter. There are a variety of tubular steam-raising converters (Fig. 7d) available, which feature radial or axial flow, with the catalyst on either shell or tube side. The near-isothermal operation of this reactor type is the most thermodynamically efficient of the types used, requiring the least catalyst volume. Lower catalyst peak temperatures also result in reduced by-product formation and longer catalyst life. 

In the mass-transfer limited region, conversion is most commonly increased by using more catalyst volume or by increasing cell density, which increases the catalytic wall area per volume of catalyst. When the temperature reaches a point where thermal oxidation begins to play a role, catalyst deactivation may become a concern. 

Both catalyst space velocity and bed geometry play a role. The gas hourly space velocity (GHSV) is used to relate the volumetric flow rate to the catalyst volume. GHSV has units of inverse hour and is defined as the volume flow rate per catalyst volume. 

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). 

Whereas changing catalyst volume or residence time rarely yields compHcations, changing temperature or pressure could iatroduce sintering. The properties of the catalyst should be measured both before and after deactivation and inlet and outlet streams should be analyzed by chromatography (qv) or spectrometry. 

The ammonia—air dilution system dilutes the vaporized ammonia by a factor of 20 to 25 with air for better admixing through the AIG and to prevent explosive ammonia—air mixtures. Once the catalyst volume is selected, the NO removal is set by the NH /NO mole ratio at the inlet of the SCR system (39). 

Performance criteria for SCR are analogous to those for other catalytic oxidation systems NO conversion, pressure drop, catalyst/system life, cost, and minimum SO2 oxidations to SO. An optimum SCR catalyst is one that meets both the pressure drop and NO conversion targets with the minimum catalyst volume. Because of the interrelationship between cell density, pressure drop, and catalyst volume, a wide range of optional catalyst cell densities are needed for optimizing SCR system performance. 

CO conversion is a function of both temperature and catalyst volume, and increases rapidly beginning at just under 100°C until it reaches a plateau at about 150°C. But, unlike NO catalysts, above 150°C there is Htde benefit to further increasing the temperature (44). Above 150°C, the CO conversion is controUed by the bulk phase gas mass transfer of CO to the honeycomb surface. That is, the catalyst is highly active, and its intrinsic CO removal rate is exceedingly greater than the actual gas transport rate (21). When the activity falls to such an extent that the conversion is no longer controUed by gas mass transfer, a decline of CO conversion occurs, and a suitable regeneration technique is needed (21). 

Southern Company Services, Inc., prime contractor for Wilsonvihe Facility. " Coal space velocity is based on settled catalyst volume. 

The tube is much longer than needed for the catalyst volume to provide a surface for preheating and to minimize temperature losses at the discharge end. The tube can be bent into a U shape and immersed in a fluidized sand bath, or it can be straight and placed inside a tubular furnace in a temperature-equalizing bronze block. Thermocouples are usually inserted... 

Actual measurement results are shown in Figure 3.4.3 Here a ROTOBERTY reactor was used with a two-stage blower pumping air at room conditions over three catalyst beds with 5, 10, and 15 cm of catalyst volume. Pressure generated was measured by a water U-tube... 

Select a volume V = 19.64 cm of catalyst to be charged both in the regular basket (case A) and in a basket (case B) that is half of the diameter of A. In cases A and B the AP and RPM are kept the same. In Cases A and C the volumetric recycle flow remains the same. In row 1 and 2 on Figure 7.2.1, if the diameter drops to half, the height must increase four times (for constant catalyst volume.) In row 4, L/dp increases four times as well. In row 5, for cases A vs B, if L/dp increases four times, has to drop to one quarter, hence u will be one half to maintain constant AP. In row 6, for u in A and B and flow cross section 1/4, the volumetric flow will be 1/8 and the recycle ratio also 1/8. In row 7, or u four times larger. 

The catalyst volume is the same on both sides. It is assumed that no diffusional rate limitation exists even in the larger pellets. That is, the chemical reaction rate is controlling. Pressure drop must be the same for both sides, so the flow has to be less over the smaller pellets to maintain the AP (L/dp)(u /2g) = constant. 

For this to be constant, if L/dp doubles then must become one half, or u to be (1/2)° = 0.71 over the smaller pellets. Since the reaction rate is the same and the catalyst volumes are the same, the number of moles converted are the same. This constant number of moles will result in the 71% flow over the smaller pellets in 1/0.71 = 1.41 times larger reactant decrease than over the larger pills. The flow volumes are relative, not absolute values. 

These are less expensive and less troublesome than tubular reactors. All the catalyst volume needed for a given conversion is usually divided in several beds or stages. In large catalyst volumes, the stages may be in separate vessels, or in small volumes in the same vessel but divided into several trays. 

The need to keep a concave temperature profile for a tubular reactor can be derived from the former multi-stage adiabatic reactor example. For this, the total catalyst volume is divided into more and more stages, keeping the flow cross-section and mass flow rate unchanged. It is not too difficult to realize that at multiple small stages and with similar small intercoolers this should become something like a cooled tubular reactor. Mathematically the requirement for a multi-stage reactor can be manipulated to a different form ... 

The total reactor volume required is independent of the number of beds in the series. This is evident because (a) all the beds operate with the same temperature profile and essentially the same pressure, (b) the inlet gas composition is the same for all the beds, and (c) the outlet gas composition is the same for all beds. Hence, the average driving force is the same for all beds, and the catalyst volume is simply related to the total production of methane. 

In summary, increasing the number of beds in series for the system shown in Figure 12 decreases the volume of recycle somewhat more than proportionately to the number of beds, decreases the recycle compressor power requirements somewhat, and has no effect on the catalyst volume. On balance, it would appear that the capital savings achieved by increasing the number of beds will be minimal. 

This refers to the total gas flow through a plane of catalyst where Nx is the mole fraction of X in the gas passing through the plane, NWeq is the mole fraction of X at equilibrium under conditions at this point in the catalyst bed, and dv is the incremental catalyst volume. 

Run LT-H catalyst volume, 0.00352 ft8, exchanger specifically installed to determine operating limits and overall ... 

Space Velocity. Most of our experimental data were developed with operation at a wet outlet space velocity of approximately 10,000 vol/vol hour. However, we do have data at space velocities of up to 25,000/hr. The pilot plant will operate at a space velocity of 5,000/hr while processing 1 million scf raw syngas/day. With operation on a once-through basis without recycle and at the indicated space velocities, catalyst volumes are minimum compared with other processes when identical over-design factors are used. 

A gas-liquid-particle process termed cold hydrogenation has been developed for this purpose. The hydrogenation is carried out in fixed-bed operation, the liquefied hydrocarbon feed trickling downwards in a hydrogen atmosphere over the solid catalyst, which may be a noble metal catalyst on an inert carrier. Typical process conditions are a temperature of 10°-20°C and a pressure of 2.5-7 atm gauge. The hourly throughput is as high as 20-kg hydrocarbon feed per liter of catalyst volume. 

It is a good idea to run the laboratory reactor without catalyst to check for homogeneous reactions. However, this method does not work when the homogeneous reaction involves reactants that do not occur in the feed but are created by a heterogeneous reaction. It then becomes important to maintain the same ratio of free volume to catalyst volume in the laboratory reactor used for intrinsic kinetic studies as in the pilot or production reactors. 

The space velocity (the gas flow divided by the catalyst volume) at which a given conversion is obtained at a certain temperature. 

The production rate can also be written as r = VSkCo, where V is the volume of the catalyst, S the Pt area per catalyst volume, k the rate constant, and Q the concentration of the reactant. 

A pulse reactor system similar to that described by Brazdll, et al( ) was used to obtain the kinetic data. The reactor was a stainless-steel U-tube, composed of a l/S" x 6 preheat zone and a 3/8" X 6 reactor zone with a maximum catalyst volume of about 5.0 cm. The reactor was Immersed In a temperature controlled molten salt bath. 

LHSV = Total Liquid Flow Rate (cc/hr) / Total Amberlyst BD20 Catalyst Volume (cc)... 

Catalyst cost constitutes 15-20% of the capital cost of an SCR unit therefore, it is essential to operate at temperatures as high as possible to maximize space velocity and thus minimize catalyst volume. At the same time, it is necessary to minimize the rate of oxidation of S02 to S03, which is more temperature sensitive than the SCR reaction. The optimum operating temperature for the SCR process using titanium and vanadium oxide catalysts is about 38CM180oC. Most installations use an economizer bypass to provide flue gas to the reactors at the desired temperature during periods when flue gas temperatures are low, such as low-load operation. 

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