Control of bed temperature

Control of bed temperature provides some special challenges. As the load is varied, either the heat transfer coefficient or immersed tube area must be varied correspondingly. This may be achieved over a moderate load change by varying the area of tubes immersed in the bed by varying the height of the bed. 

Catalytic oxidation of naphthalene to phthalic anhydride is the seccxid application of FCB which was initiated in 1945. Riley (R12) reported problems on afterburning and on the control of bed temperature, and introduced many improvements in the fluidized bed. However, the most important change to make scale-up easier and to achieve good performance was the use of microspherical catalyst, or moderately active German-type catalyst (Bll, G14). 

Relatively low uniform temperature allows good control of bed temperature to be achieved. 

The transfer of heat between fluidized solids, gas and internal surfaces of equipment is very good. This makes for uniform temperatures and ease of control of bed temperature. 

If air (or oxygen) and steam are both passed through a high-temperature bed of coal or coke these reactions can be balanced, thus controlling the bed temperature and the fusion of the ash. In the higher pressure Lurgi process the gas obtained is high in methane, formed in reactions such as... 

Catalytic activity was determined with a fixed bed microreactor which consisted of two coassial quartz tubes (i.d. 35 and 16 mm) to allow feed gas preheating and heated in an electrical oven (Watlow) with a temperature controller. The bed temperature was monitored by A K-type thermocoupling. On-line analysers (ABB) for CO, C02, CH4, ... 

In the same manner as in hydrocracking (Dolbear, 1997), hydrogen is added at intermediate points in hydrodesulfurization reactors. This is important for control of reactor temperatures. The mechanical devices in the reactor, called reactor internals, which accomplish this step are very important to successful processes. If redistribution is not efficient, some areas of the catalyst bed will have more contact with the feedstock. This can lead to three levels of problems ... 

The U.O.P. Chamber-type unit contains the catalyst in a reactor in which the catalyst is separated into a number of beds. Temperature control is accomplished by quenching between the beds with cold liquid effluent and by recycling spent propane and butane effluent into the reactor feed. This is not a heat-balance operation and a preheater using an outside source of heat is required in addition to the feed-to-products exchanger. A depropanizer and debutanizer are usually used to fractionate the effluent. Figure 2 shows a flow diagram of a chamber type of polymerization unit. 

In grate-fired boilers, the biomass is fed in a thin layer, so it is evenly distributed over a sloped stationary, traveling, or vibrating grate. Improved control over the combustion process can be achieved with better carbon conversion. Fluidized-bed combustors are more complex systems, but offer much better control of combustion temperature, improved carbon conversion, and fuel flexibility. Using a boiler to produce both heat and electricity (co-generation) can improve the overall system efficiency to as much as 85 percent. Boiler efficiencies are affected by fuel moisture content, air-fuel ratio, excess air, combustion temperature, and biomass ash content. 

Fluidized-bed reactors have proposed as alternatives to circumvent the problems mentioned above (ALMA process, Alusuisse, Lummus). These allow better control of the temperature and higher feed concentrations. In addition the reactor system is less prone to explosions. A major problem in this reactor is attrition of the catalyst. In recent literature this problem has received a great deal of attention. A second drawback of the fluidized-bed reactor is the intrinsic mixing behaviour, which lowers the conversion. 

Most of the reactions that proceed in FFB are strongly exothermic or endothermic, and heat must be removed from or supplied into the bed in order to control the bed temperature. So heat transfer in FFB is of great importance for design and operation. 

These techniques operate in accordance with similar principles (Fig. 3.7). The preheated feed, normally heated by heat exchange with the reactor effluent, is introduced in the reactor in a downflow stream. Conversion takes place m the presence of add catalysts such as phosphoric add deposited on silica or siUco aluminates. placed in several fixed beds. This arrangement allows eflecdve control of the temperature rise due to the exothermic nature of the reaction, by the injection between the beds of. a quench liquid consisting of unconvened C4. The operating conditions are moderate (temperatures average 120°C, limits 50 to 200 C, pressure I to 4. tO a absolute). The reactor effluent... 

This conversion is conducted at moderate temperature and pressure (lOO C, IS. 10 Pa absolute), and possibly in the presence of a hydrocarbon diluent, for better control of the temperature rise in the catalyst beds, due to the high exothennicity of the reaction, which is itself related to the high diolefmic content of the initial Cj cut As a rule, the feed is introduced in a downflow stream into the reactor, which contains several beds of a noble metal catalyst on alumina. Quench by recycling and diluent injection is carried out between the beds. The diluent is recovered, by distillation in a depentanizer, after flash to eiimioate the inert compounds introduced with hydrogen gas at the same time as the feedstock. The leading licensors include FP and Shett, etc... 

To provide an illustration, the flow sheet of the IFP process shown in Fig. 3.12 comprises two possible variants. The simpler corresponds to the direct use of the etherified solution in the gasoline pool, without separating e excess methanol contained. Operations are conducted with two reactors in series the first with an upflow stream and expanded bed with recycle of part of the previously cooled effluent for better control of the temperature rise, and the second with a downflow stream and a fixed bed. The more complex involves the recovery of excess methanol, first by azeotropic distillation in a depentanizer with part of the unconverted hydrocarbons, and then by water washing of this raffinate. The hydrocarbon phase is added to the bottom of the depentanizer. The water/methanol mixture is distilled to recover and recycle the alcohol to the etherification staee. 

Between the catalyst beds or in the axis of the shell cooling is often provided by means of hydrogen rec cle. for example, for better control of the temperature rise, which must not be more than 50 to 75 C between the reactor inlet and exit. The reaction section effluent is cooled, sometimes quenched by a recycle, and sent to a high pressure separator where the hydrogen is recovered. This is recycled after passage through a toluene absorber, to rid it of traces of entrained aromatics, after purification or purge. 

Control of the temperature throughout the reforming catalyst bed can be established by use of a monolithic catalyst. The heat transfer control can be accomplished by combining three effects that monolithic catalyst beds can impact significantly (1) direct, uniform contact of the catalyst bed with the reactor wall will enhance conductive heat transfer (2) uniformity of catalyst availability to the reactants over the length of the flow will provide continuity of reaction and (3) coordination of void-to-catalyst ratio with respect to the rate of reaction will moderate gas-phase cracking relative to catalytically enhanced hydrocarbon-steam reactions. This combination provides conditions for a more uniform reaction over the catalyst bed length. 

The plant was operated between 650°-750°C in the pyrolysis temperature under stable controls of pressure, temperature and sand circulation. Since the fluidized bed has an effect of heat accumulation, it is easy to control the operating temperature, though the heterogeneous refuse is fed to the bed. Consequently, no damage of reactor materials due to excessive temperature were recognized. 

A primary example is the resolution of optical isomers by continuous crystallization in fluid beds. Control of low supersaturation by control of the temperature difference between the continuous feed and the seed bed is critical to maintaining an essentially all-growth regime in which the individual isomers grow on their respective seeds in separate crystalfi-zers. The seed beds in both crystallizers are massive in relation to the amount of racemic solution passing through in order to present sufficient seed area to maintain low supersaturation. Uncrystallized isomers in the overhead streams are recycled to dissolve additional racemic feed. Crystal size is maintained by sonication. See Examples 7-6 and L1-6 for a discussion of resolution of optical isomers by continuous crystallization. 

This preheated feed is introduced at the top of a reactor with superimposed catalyst beds, formed of phosphoric add deposited on kielselguhr, which may be promoted by boron trifluoride (t/OP), at a temperature between 190 and 200°C and a pressure between 35 and 4.10 Pa absolute The LHSV (Liquid Hourly Space Velotity) in relation to benzene is about 1.5. Propane, containing a calculated amount of water (about 300 ppm in relation to the reactor feed) is injected between the catalyst beds to remove the heat liberated by the reaction, for better control of the temperature rise, which must not exceed 30 to 50°C in each of the beds, and to maintain the catalyst system in a suitable state of hydration. Performance is accordingly improved. 

Among the Ni/Mg/Al samples, the microwave-treated SAl catalyst is more active and selective towards syngas than the merely stirred catalyst BOl. However, under hard reaction conditions differences are less important. In fact, with very active and selective catalysts, differences can be smoothed by a temperature increase, which displaces the reaction towards the equilibrium. When using diluted mixtures the heat produced in the exothermic reaction is reduced, thus achieving a better control of the temperature inside the catalytic bed. For this reason the beneficial effect of the MWHT aging treatment was more notable in the tests carried out at 500°C. 

The present work has established the potential of a laboratory multi-layer-packed bed reactor for xylanase production in solid-state fermentation by P. canescens. Optimal production obtained is largely superior to those related in literature. Two problems are identified. Initially, the forced aeration causes sporulation of the Penicillium strain and so decreases xylanase production. In addition, increasing the cultural surface more than 100 g/layer of soya oil cake (1-cm bed height) decreases significantly xylanase production and constitutes a disadvantage for this process. To increase efficiency of the multi-layer-packed bed reactor, we can increase the dimensions of the layers. Why not associate mixing the culture medium, moistened air, and a rigorous control of the temperature at the inner the reactor Another approach would be to develop a reactor in this direction to improve the production. 

The catalyst was Rh on y-alumina in the form of particles with diameter of 0.1-0.3 mm. In some measurements, supported Pt or Pd were also utilized. The catalyst was packed in a specially designed, 8-mm i.d. tubular reactor, between two layers of inert material. A forced ventilation oven allowed control of reactor temperature. The expressions reactor inlet and reactor outlet adopted in the text to describe the reaction front motion, are referred to the catalyst bed only, without taking into account the inert layer. 

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