WO1999016543A1

WO1999016543A1 - Endothermic process

Info

Publication number: WO1999016543A1
Application number: PCT/GB1998/002891
Authority: WO
Inventors: Samuel David Jackson; Frank King; David Graham Shipley; Edmund Hugh Stitt
Original assignee: Imperial Chemical Industries Plc
Priority date: 1997-10-01
Filing date: 1998-09-24
Publication date: 1999-04-08
Also published as: AU9177398A; GB9720730D0; WO1999016543A8

Abstract

An endothermic reaction is effected using a heat transfer bed, containing a variable oxidation state material, which is cycled through a succession of stages including an oxidative regeneration stage employing passage of a hot oxidising gas through the bed, a reductive regeneration stage employing passage of a reducing gas that is oxidised with the evolution of heat through the bed, each cycle of the process having a duration of less than 5 minutes. The convertible material comprised a chromic containing catalyst.

Description

Endothermic process

This invention relates to an endothermic process wherein a reactants stream is passed through a heated body, which may comprise a bed of material exhibiting catalytic activity for the desired reaction. Where used, such a catalyst material may be supported on a suitable substrate. In such a process, the heated body, hereinafter termed a heat transfer bed, provides the heat required for the endothermic reaction and it is periodically necessary to reheat the bed. Such re-heating is herein termed regeneration and may also effect, or accompany, removal of by-products (by which term we include the products of side reactions), which impair the activity of any catalytic material.

The re-heating may effect conversion of catalytic material from an inactive state to the active state. As an example of such a process, there may be mentioned the dehydrogenation of paraffins, e.g. propane, to form olefins, e.g. propene. An example of such a process is described in "Oil &

Gas Journal" December 8, 1980 pages 96-101. In that process propane is passed through a hot adiabatic catalyst bed in which the catalyst material promotes the dehydrogenation, giving propene and hydrogen. The dehydrogenation reaction is endothermic, resulting in cooling of the catalyst bed. As a result of thermal cracking side-reactions, carbon is liable to be deposited. The bed is then regenerated by passing heated air through the catalyst: this also serves to burn off deposited carbon.

Such reactions may be effected using a plurality of static beds which are cycled, by switching the flows of reactants etc., through reaction and regeneration stages. Alternatively, and preferably, a rotatable heat transfer bed is employed which is rotated through a regeneration zone, then through a reaction zone, and then returned to the regeneration zone.

US-A-2 704 741 describes a reactor for that type of process wherein a catalyst bed is disposed in compartments in a rotating annular vessel: the vessel is disposed between, and sealed against, static outer and inner vessels which are divided into compartments. Provision is made for supply of reactants and regeneration gas to appropriate compartments and for receipt of products and spent regeneration gas from other compartments. The reactants flow radially through the annular vessel between the appropriate compartments of the inner and outer vessels.

US-A-4 418 046 describes a similar arrangement wherein, instead of the catalyst bed being disposed in separate compartments in the rotating body, the catalyst bed is in the form of a honeycomb structure so that the honeycomb wails serve to separate adjacent flow passages: in this reference, depending on the honeycomb configuration, the flow can be radial or axial.

EP-A-0 193 511 also describes a similar arrangement wherein a honeycomb catalyst bed is employed with the honeycomb having cells providing axial passages and the flow of reactants etc. is axial. Endothermic reactions, such as propane dehydrogenation, are favoured thermodynamically by operation at high temperatures. For example propane dehydrogenation is normally effected at temperatures of the order of 500-700°C. In the regeneration stage, an oxidising gas, e.g. air, is passed through the catalyst bed. Not only does the regeneration gas serve to burn-off by-products such as deposited carbon, but also supplies the heat for the endothermic reaction. Thus hot regeneration gas heats the catalyst bed so that transfer of the catalyst bed from the regeneration 5 stage into the reaction stage transfers heat to the reaction stage.

Heretofore, unless the regeneration stage occupies a very large proportion of the cycle, thus rendering the size of catalyst bed employed large and uneconomic, it has been necessary to employ very hot gas for the regeneration and/or use very high gas mass flow rates in order to transfer sufficient heat in a reasonable time. The use of very hot gases leads to problems of o thermal stressing of the catalyst bed and/or undesired reactions, while the use of very high gas mass flow rates necessitates that the regeneration pressure is substantially greater than the reaction pressure. That in turn leads to engineering difficulties, particularly in a rotary system of the type described in the above references.

It is known that if the heat transfer bed contains a material having a first oxidation state during 5 the reaction and that can be oxidised to a higher oxidation state, oxidation of that material during an oxidative regeneration stage to the higher oxidation state, followed by a reductive regeneration stage wherein the material is reduced from that higher oxidation state to the lower oxidation state, can result in sufficient heating of the heat transfer bed to provide a significant part, if not all, of the necessary reaction heat. As a result the use of regeneration gas at very high temperatures and/or 0 at pressures substantially greater than the reaction pressure is no longer necessary. The material, hereinafter termed convertible material, having the higher and lower oxidation states preferably, but not necessarily, has catalytic activity for the desired reaction: in such cases, the catalytically active state will be that of the lower oxidation state. Examples of such processes are described in GB-A-579 477 and GB-A-837 707. Thus GB-A-579 477 describes dehydrogenation of 5 hydrocarbons using a cycle having an oxidative regeneration step employing a catalyst such as chromia supported on alumina. Typical process cycles described include a step of reduction of the catalyst with hydrogen after the oxidative regeneration. In order to avoid "run-away" reactions, the catalyst temperature is maintained at a temperature at which adiabatic balance is achieved and the shortfall in reaction heat is supplied by sensible heat from the reactants: thus the reactants are 0 supplied at a temperature above the average catalyst temperature. GB-A-837 707 describes dehydrogenation of hydrocarbons employing a regenerable chromia catalyst wherein part of the chromia is oxidised to the hexavalent state during the oxidative regeneration process. The description indicates that the heat of combustion of the by-product carbon during the regeneration step can supply the heat required for the dehydrogenation reaction and that the reduction of the 5 hexavalent chromium compound, which reduction occurs during the reaction stage, can supplement the heat. In these prior art processes, the cycle times are relatively long, the reaction stage typically being at least 10 minutes. We have found that in this type of process, the cycle time is desirably much shorter, less than 5 minutes, for several reasons. Firstly a short cycle time implies that the reaction time during each cycle is short; this is advantageous since the activity of the catalyst 5 declines rapidly and secondly, if the reaction time is short, the by-product carbon is deposited in a form which renders it readily removed without employing long regeneration times. Unlike the aforesaid processes using long reaction times, the amount of carbon deposited during each reaction stage is relatively small and so the combustion thereof during the subsequent oxidative regeneration stage only supplies a small part of the heat required for effecting the reaction. A o further, important, advantage is that because of the short cycle time, a significant proportion of the heat required for the reaction may be provided by the oxidation and reduction of the catalyst.

Accordingly the present invention provides a process for performing an endothermic reaction comprising cycling a heat transfer bed, containing a solid convertible material that can be oxidised from a lower oxidation state at which the reaction occurs to a higher oxidation state, through a 5 succession of stages: a) an oxidative regeneration stage including a step wherein an oxidising gas is passed through the heat transfer bed, whereby said convertible material is oxidised from said lower oxidation state to said higher oxidation state, and any deposited by-products of said reaction are oxidatively removed, with consequent heating of the heat transfer bed; 0 b) a reductive regeneration stage wherein a reducing gas that is oxidised with the evolution of heat is passed through said heat transfer bed, whereby said convertible material is reduced from said higher oxidation state to said lower oxidation state with consequent further heating of the heat transfer bed; and c) a reaction stage wherein feedstock is passed through the heated heat transfer bed, whereby 5 said feedstock is reacted to give a products stream and said heat transfer bed is cooled, each cycle of the process having a duration of less than 5 minutes.

Thus in the present invention a substantial part, preferably at least 60%, of the heat required for the reaction is provided by the heat evolved by chemical reactions, i.e. oxidation of the convertible material, oxidation of the reducing gas by the convertible material being reduced from its 0 higher oxidation state to its lower oxidation state, and oxidation of any deposited by-products. As a result only a minor part, if any, of the heat required for the reaction need be provided by the sensible heats of the oxidising and reducing gases. Indeed, in some cases there may be a net gain in sensible heat by the oxidising and/or reducing gas.

The oxidising gas employed is preferably air, but in some cases it may be desirable to use a 5 gas having a lower oxygen content than air. Thus air, or oxygen diluted with an inert diluent, e.g. nitrogen, may be used. The reducing gas is preferably hydrogen. However in some cases, where the feedstock can effect reduction of the convertible material to the lower oxidation state, the reduction and reaction stages can be combined by using the feedstock as the reducing gas: thus in the initial part of the reaction stage some of the feedstock effects reduction of the convertible material from its higher oxidation state to the lower oxidation state and then the desired endothermic reaction of the remaining feedstock proceeds.

The invention is further described by reference to catalytic dehydrogenation of propane but it will be appreciated that by suitable choice of the convertible material and conditions, other endothermic reactions, for example thermal cracking of hydrocarbon feedstocks such as ethane and higher hydrocarbons to olefins, may be effected.

Catalytic propane dehydrogenation, for example at atmospheric pressure and at a temperature of the order of 600°C, proceeds only partially to completion and produces some by-products. A typical products stream has the approximate volume composition: propane 50% hydrogen 24% propene 22% methane 3% ethane 1 % ethene <0.5% and about 0.5 kg of carbon are deposited on the catalyst for every kmol of propane fed to the heat transfer bed containing the catalyst. The heat required for the reaction to produce this products stream is about 38 MJ/kmol of propane feed. This heat is supplied from the heat transfer bed. Where the products stream has a temperature above the feedstock propane temperature, as will normally be the case, the heat transfer bed also has to supply the additional sensible heat in the products stream. During the oxidative regeneration stage, combustion of the deposited carbon can provide about 16.6 MJ per kmol of propane fed to the dehydrogenation stage, thus leaving about 21.4 MJ per kmol of propane fed to the dehydrogenation stage, together with any required sensible heat, to be supplied from another source. In the present invention, part, if not all, of this additional heat is supplied by the reductive regeneration stage. Laboratory studies have shown that using a chromia containing catalyst material for propane dehydrogenation, the oxygen to chromium atomic ratio in the active catalyst is about 1.3 or less. In the oxidative regeneration stage, in addition to burning off the deposited carbon, the catalyst material is oxidised by means of the oxidising gas to an oxygen to chromium atomic ratio of 1.5. During the reductive regeneration step, the oxidation state of the catalyst is reduced back to an oxygen to chromium atomic ratio of about 1.3 with the released oxygen oxidising the reducing gas with the evolution of heat. Hence about 0.2 atoms of oxygen are released per atom of chromium in the catalyst. Since hydrogen is produced in the propane dehydrogenation reaction, it is possible to use some of the hydrogen produced as reducing gas: thus the hydrogen may be separated from the products or part of the products recycled as reducing gas. Where hydrogen is used as reducing gas, the oxidation releases 242 MJ of heat per kmol of hydrogen, i.e. 0.93 MJ of heat per kg of 5 chromium. Hence to supply the 21.4 MJ of additional heat per kmol of propane fed to the dehydrogenation stage, about 23 kg of chromium has to be subjected to the oxidation/reduction cycle.

Design calculations on a preferred arrangement using a rotating catalyst bed as described hereinafter show that about 33 m³ of a honeycomb chromia/alumina catalyst containing 5% by o weight Cr and having an open area of 58% can effect the dehydrogenation of propane at 600°C fed at a rate of 1 kmol/s. The 33 m³ of honeycomb catalyst contains about 1800 kg of chromium, and so one cycle of the catalyst bed through the oxidative and reductive regeneration stages releases about 1674 MJ of heat. Hence for this released heat to provide all of the required 21.4 MJ of additional heat per kmol of propane, the catalyst needs to be subjected to the cycle of the oxidative 5 and reductive regeneration steps once every 1674/21.4 seconds, i.e. approximately every 78 seconds. In such a case the reaction time is typically about 1 minute and thus constitutes about 75% of the cycle time. If the cycle time is increased to 150 seconds, it is seen that only about 65% of the required additional 21.4 MJ of heat per kmol of propane is supplied by the oxidation/ reduction of the catalyst. 0 Where hydrogen is used as the reducing gas, approximately 0.1 kmol of hydrogen is required to produce the reaction heat per kmol of propane fed to the dehydrogenation stage: since the dehydrogenation process produces about 0.3 kmol of hydrogen per kmol of propane fed to the dehydrogenation stage, the net amount of hydrogen produced is about 0.2 kmol.

Instead of using hydrogen, e.g. part of the products stream, as the reducing gas, propane may 5 be employed. Thus, it has been found in laboratory studies that the propane is oxidised by the catalyst material approximately in accordance with the equation

2C₃H₈ + 6[0] > 3C + 3C0₂ + 8H₂ where [O] represents oxygen supplied by the chromia in the catalyst material. The carbon is deposited in addition to the 0.5 kg (per kmol of propane fed to the dehyrogenation stage) deposited 0 as a result of the side reactions ocurring during the dehydrogenation of propane. The carbon deposited during the reduction of the catalyst material is burnt off in the next oxidative regeneration stage and so the net oxidation/reduction reaction is

C₃H₈ + 30₂ > 3C0₂ + 4H₂ which releases about 1077 MJ of heat per kmol of propane utilised in the reduction step. Hence, 5 per kmol of propane fed to the dehydrogenation stage, the additional heat can be supplied by using about 0.02 kmol of propane as the reducing gas in the reductive regeneration stage. If there is an increase in sensible heat of the products stream, the necessary heat may be provided by decreasing the cycle time, so that a greater amount of heat is provided by the oxidation/reduction of the convertible material. Alternatively a fuel, e.g. propane and/or hydrogen, may be added to the feed to the oxidative regeneration step so that the heat produced by combustion therein of any deposited carbon is augmented by combustion of the fuel. As an alternative to chromia, a supported platinum catalyst may be employed. In the present invention, the convertible material thus acts as an oxygen carrier from the oxidative regeneration stage to the reductive regeneration stage and the bed containing the convertible material acts as a heat carrier. In the present invention, the cycle time should be relatively short, less than 5 minutes, and preferably less than about 3 minutes, and particularly 0.5 to 2.5 minutes. The reaction time, i.e. the time the heat transfer bed is in the reaction stage, is preferably at least 40% of the cycle time, and preferably is 50-80% of the cycle time. As a result of short reaction duty periods, deposited products, e.g. carbon, of side reactions do not have time to age to any significant extent: such ageing renders removal of the deposits more difficult in the subsequent oxidative regeneration stage thus increasing the time needed for that oxidative regeneration step. Increasing the oxidative regeneration time increases the cycle time and hence increases the size of the catalyst bed needed to process the feedstock at a given rate. In this regard it is noted that in the aforesaid GB-A-837 707 the reaction and regeneration times are relatively long. That reference indicates that relatively long oxidative regeneration times were necessary in order to oxidise the chromium in the catalyst to the hexavalent state. Laboratory studies have shown that at the relatively short reaction and regeneration times contemplated in the present invention, hexavalent chromium is not formed.

Since the desired reaction is generally reversible and favoured by high temperatures, it is desirable, in order to minimise the extent to which the reverse reaction occurs, to cool the products rapidly from the reaction temperature to a temperature below that at which the reverse reaction occurs at a significant rate. In the present invention heating the feedstock to the reaction temperature and rapid cooling of the products may be effected in the heat transfer bed, so that the latter acts as a heat exchange medium. In this way the increase in sensible heat of the products stream can be minimised, so that the amount of reducing gas required to be oxidised to provide heat can be minimised.

While the process may be operated with a single heat transfer bed with intermittent flow of oxidising gas, reducing gas, and feedstock as the heat transfer bed is cycled through the various stages, it is preferred that either a plurality of beds are employed so that one or more beds are undergoing regeneration while one or more other beds are undergoing the reaction stage, or, more preferably, a single heat transfer bed is employed which is moved, e.g. by rotation, through zones in which the various stages are taking place. Thus a preferred form of the invention provides a process for performing an endothermic reaction using a moving heat transfer bed containing a solid convertible material that can be oxidised from a lower oxidation state at which the reaction occurs to a higher oxidation state, said moving bed being in the form of a rotatable member having a multiplicity of through flow passages, the walls of which are formed from, or are coated with, or which passages contain said convertible material, said process comprising: a) continuously rotating said rotatable member about its axis, with each rotation having a duration of less than 5 minutes, whereby each flow passage is moved through a succession of zones including an oxidative regeneration zone, a reductive regeneration zone, a reaction zone, and then returned to the oxidative regeneration zone; b) passing an oxidising gas stream through said oxidative regeneration zone, whereby the convertible material that is in said oxidative regeneration zone is oxidised from said lower oxidation state to said higher oxidation state, and any deposited by-products of said reaction are oxidatively removed, with consequent heating of that portion of the heat transfer bed that is in said oxidative regeneration zone; c) passing a reducing gas that is oxidised with the evolution of heat through said reductive regeneration zone, whereby the convertible material that is in said reductive regeneration zone is reduced from said higher oxidation state to said lower oxidation state with consequent further heating of that portion of the heat transfer bed that is in the reductive regeneration zone; and d) passing feedstock through said reaction zone, whereby said feedstock is reacted to give a products stream and that portion of the heat transfer bed that is in said reaction zone is cooled.

As mentioned above, in some cases the feedstock may be used as the reducing gas. Accordingly a preferred form of the invention provides a process wherein a feedstock is subjected to an endothermic reaction using a moving heat transfer bed containing a solid convertible material that can be oxidised from a lower oxidation state at which the reaction occurs to a higher oxidation state, and reduced to the lower oxidation state with the evolution of heat by said feedstock, said moving catalyst bed being in the form of a rotatable member having a multiplicity of through flow passages, the wails of which are formed from, or are coated with, or which passages contain said convertible material, said process comprising: a) continuously rotating said rotatable member about its axis, with each rotation having a duration of less than 5 minutes, whereby each flow passage is moved through a succession of zones including an oxidative regeneration zone, a reaction zone, and then returned to the oxidative regeneration zone; b) passing an oxidising gas stream through said oxidative regeneration zone, whereby the convertible material that is in said regeneration zone is oxidised from said reactive state to said higher oxidation state, and any deposited by-products of said reaction are oxidatively removed, with consequent heating of that portion of the heat transfer bed that is in said oxidative regeneration zone; and c) passing said feedstock through said reaction zone, whereby the convertible material in the initial part of said reaction zone is reduced from said higher oxidation state to said lower oxidation state with consequent further heating of that part of the heat transfer bed in that initial part of the reaction zone, and, in the remaining part of the reaction zone, said feedstock is reacted to give a products stream and that part of the heat transfer bed in said remaining part of the reaction zone is cooled. The reaction stage may be effected in two parts with the heat transfer bed effecting heat exchange between the feedstock and the products. In the first part of the reaction stage, the feedstock is passed through the heat transfer bed in a first direction and then, in a second part, the partially reacted feedstock from the first part of the reaction stage is passed through the heat transfer bed in the opposite direction. Where the feedstock acts as the reducing gas, the reduction of the convertible material will normally occur in the first part of the reaction stage. During the regeneration stage, the oxidising gas, and reducing gas, if a separate reductive regeneration step is employed, are preferably passed through the heat transfer bed in the direction opposite to that employed in the first part of the reaction stage.

In such a process employing two reaction stages, where the convertible material is a catalyst for the reaction, while all the heat transfer bed may comprise active catalyst, or a support carrying such a catalyst, in some cases it may be desirable to have the portion of the heat transfer bed at the end into which the feedstock is fed during the first reaction stage, free from catalytically active material: in this way the initial part of the heating of the reactants, and the final part of the cooling of the products, taking place in the bed is effected in the absence of catalyst, thereby minimising the extent to which undesired reactions, e.g. the reverse reaction, take place.

The invention is illustrated by reference to the accompanying drawings wherein: Figure 1 is a diagrammatic representation of a propane dehydrogenation process employing separate oxidation and reduction stages and a single reaction stage;

Figure 2 is a diagrammatic representation of a propane dehydrogenation process employing separate oxidation and reduction stages and two reaction stages; Figure 3 is a graphical representation of the temperature profiles at various stages in the process of Figure 2; Figure 4 is a diagrammatic representation of a rotatable catalyst bed arrangement for use in the process of the embodiment of Figure 2. In the processes of Figures 1 and 2 and the convertible material is a catalyst for the reaction in the form of the convertible material, e.g. a chromia composition, supported on an alumina honeycomb.

Referring to Figure 1 , the oxidative and reductive regeneration stages are designated A and B respectively, and the reaction stage is designated C. A post-reaction hydrogen flushing stage is designated D and purge stages before and after the oxidative regeneration stage are designated E1 and E2 respectively. The catalyst bed is designated generally by reference 10 and has first and second ends 12 and 14 respectively. The alumina honeycomb passages extend longitudinally from the first end 12 to the second end 14. The bed 10 is cycled through the stages in the sequence: purge E1 , oxidative regeneration A, purge E2, reduction regeneration stage B, reaction stage C, and hydrogen flush stage D before returning to purge E1.

Nitrogen is fed from a supply via line 16 and heated in a heat exchanger 18 by means of a hot gas stream 20 from e.g. a fired heater (not shown). In the first purge stage E1 part of the heated nitrogen stream is fed via line 22 to the first end 12 of the catalyst bed 10. This serves to sweep from the body 10 any gases remaining from the previous cycle. The spent purge gas is fed via line 24 to a second heat exchanger 26 and then is vented, via line 28. In the oxidative regeneration stage A, air, supplied via line 30, is mixed with part of the nitrogen stream supplied by line 16 and then heated in heat exchanger 26 by heat exchange with the hot gas from line 24. The hot air/nitrogen mixture is then fed via line 32 to the first end 12 of the catalyst body. During the oxidative regeneration stage A deposited coke is burnt off and the catalyst oxidised to its higher oxidation state. Both these reactions are exothermic and so the catalyst body 10 is heated. The spent air/nitrogen mixture from the oxidative regeneration stage A is added, via line 34, to the spent purge gas from the first purge stage E1 and so augments the hot gas fed to heat exchanger 26 via line 24. Instead of using a heated air/nitrogen mixture as the oxidising gas in the oxidative regeneration stage A, the oxidising gas may be the hot combustion products of burning a suitable fuel in an excess of air.

The catalyst body 10 is then submitted to a second purge stage E2 wherein the remainder of the heated nitrogen from heat exchanger 18 is passed, via line 36, to the first end 12 of the catalyst body. This purge step E2 serves to remove any air remaining in the catalyst body. The spent purge gas from the second end 14 of the catalyst body 10 is then fed, via line 38, to the spent purge gas from the first purge stage E1 and to the spent air/nitrogen mixture from the oxidative regeneration stage A thus augmenting the gas fed to heat exchanger 26 via line 24.

In the next stage, the reductive regeneration stage B, hydrogen is fed via line 40 to a third heat exchanger 42 and then, via line 44, to the first end 12 of the catalyst body 10. The convertible material is thus reduced to its lower oxidation state. This reduction is exothermic and so the catalyst body becomes further heated. The spent hydrogen stream leaves the second end 14 of the catalyst body via line 46. In the reaction stage C, propane is fed, via line 48, to a fourth heat exchanger 50 wherein it is heated and then fed, via line 52, to the second end 14 of the catalyst body 10 so that it flows counter-current to the flow of gas in the purge and regeneration stages. During the reaction stage C, the propane undergoes endothermic dehydrogenation, forming propylene and hydrogen, with the heat of reaction being supplied by the hot catalyst body 10. The reaction products are fed, via line 54, from the first end 12 of the catalyst body to heat exchanger 50 where the reaction products cool by heat exchange with the propane feed. The cooled reaction products are fed, via line 56, to a separation stage (not shown). After the reaction stage C, the spent hydrogen from the reductive regeneration stage B is fed, during a hydrogen flushing stage D, via line 46 to the second end 14 of the catalyst body 10 so that the stream of spent hydrogen flushes any residual propane and dehydrogenation products from the catalyst body 10. Thus, where the catalyst body 10 is in the form of a rotating honeycomb, some of the propane/reaction products will be trapped in the honeycomb passages as they move from the reaction stage B to the hydrogen flushing stage D. The hydrogen flushing stage D serves to effect recovery of this "transported" propane/reaction products. The spent flushing stream is then fed, from the first end 12 of the catalyst body 10, via line 58, to heat exchanger 42, where it serves to heat the hydrogen stream employed for the reductive regeneration stage B. After passage through heat exchanger 42, the spent flushing stream is added to the cooled products stream 56 and fed to the separation stage. The cycle is now complete so that the catalyst body can be subjected to the purge stage E1 of the next cycle.

The purge stages E1 and E2 serve to separate the oxygen-containing stream from the hydrogen-containing stream and thus minimise explosion hazards.

The process of the embodiment of Figures 2 and 3 is similar to that of Figure 1 but employs two reaction stages which are designated C1 and C2. For simplicity, in Figure 2, the heat exchangers, the purge steps before and after the oxidative regeneration stage A, and the hydrogen flushing stage D are not shown. It is assumed for simplicity that purge steps and the hydrogen flushing step have a negligible effect upon the temperature profiles. Thus it is assumed for simplicity that the temperatures of the catalyst bed at the start of the oxidative regeneration stage A are the same as those at the end of the second reaction stage C2 and those at the start of the reductive regeneration stage B are the same as those at the end of the oxidative regeneration stage A. It is assumed that the reaction proceeds at a significant rate only at temperatures above a temperature designated TO.

At the start of the oxidative regeneration stage A, the catalyst bed at the first end 12 has a temperature T1 while that at the second end 14 has a lower temperature T2. Temperature T2 is preferably below temperature TO; however, as described below, this is not necessarily the case. At the finish of the reductive regeneration stage B, the first end 12 of the catalyst bed has had its temperature increased from T1 to T3 while the second end has had its temperature increased from T2 to T4. The temperature of portions of the catalyst bed between the ends will generally be between T3 and T4 although, as a result of heating by heat evolved during regeneration, there may be intermediate portions of the catalyst bed having a temperature above T4. Temperature T4 will generally be somewhat below T3, but will normally be above TO.

At the start of the first reaction stage C1, it is thus assumed that end 12 of the catalyst bed 10 has a temperature T3 while end 14 has a temperature T4. During the first reaction stage, feedstock propane at a temperature T5, below T4 and well below TO, is fed via line 52 to end 14 of the catalyst bed and so is heated by heat exchange from the catalyst bed towards T3 and the desired reaction partially takes place with the heat required for the reaction being supplied by heat in the catalyst bed. As a result of the heat exchange occurring during the first reaction stage, at the finish of the first reaction stage C1, the temperature of the end 14 of the catalyst bed has fallen to a temperature T6 somewhat above T5. The feedstock inlet temperature T5 is preferably, but, as explained below, not necessarily, sufficiently below TO that T6 is still well below TO. Likewise the temperature of end 12 of the bed has fallen from T3 to a temperature T7 which however is still above the temperature TO.

During the second reaction stage C2, the partially reacted feedstock that has left the end 12 of the catalyst bed is fed back through the catalyst bed from end 12 to end 14. Further reaction of the feedstock takes place giving a products stream which leaves end 14 of the catalyst bed via line 54a. Since, during the second reaction stage C2, the temperature at the end 14 of the catalyst bed is substantially below the temperature TO, cooling of the products stream takes place by heat exchange with the catalyst bed. This heat exchange results in end 14 of the catalyst bed being heated during the second reaction stage from T6 to T2. At the start of the second reaction stage C2 the end 14 is at temperature T6, which, as indicated above, is determined by the feedstock inlet temperature. Although the products leaving the second reaction stage C2 via line 54a will have a temperature above that of end 14, and end 14 will be heated during the second reaction stage to T2, provided that T2 is sufficiently below TO and/or, as explained below, the products leaving the end 14 at the finish of the second reaction stage C2 are mixed with products leaving end 14 at earlier periods of the second reaction stage, the products leaving the bed via line 54a during the second reaction stage C2 can be cooled to a temperature below TO. Also, as a result of the further reaction and heat exchange, the temperature of the catalyst bed at the end 12 falls during the second reaction stage from T7 to T1. The bed is then ready for the start of the next cycle.

Where the process is operated in a continuous manner, for example where, as described above, a rotatable member carrying the catalyst bed is rotated to continuously move flow passages, the walls of which are formed from, or are coated with, or which passages contain, catalyst material, through oxidative regeneration and the first and second reaction zones in succession, it will be appreciated that the product stream will be a mixture of products leaving flow passages that have been in the second reaction zone for varying periods and so will have a temperature corresponding to the average of the products outlet temperature at the various times of the second reaction stage. Thus the effective products outlet temperature will be lower than that of products leaving flow passages nearing the end of the second reaction stage. Provided that the temperature of the products mixture is below TO, it may not be necessary that T2 is below TO.

Likewise, as shown by the broken line 60, the partially reacted feedstock leaving the first reaction zone at various periods thereof may be mixed before passing to the second reaction zone. Figure 2 graphically shows the temperature profiles: for simplicity the profiles are shown as straight lines whereas in reality they will be curves: thus line I shows the profile of the end 12 of the catalyst bed; line II shows the profile of the end 14 of the catalyst bed; line III indicates the temperature T5 of the feedstock fed via line 52; line IV shows the profile of the products stream leaving end 14 assuming a constant difference in temperature between the products stream and the end 14; line V indicates the temperature Tav if the products stream from the various periods of the second reaction stage are mixed; and line VI indicates the temperature TO below which the reaction does not occur at a significant rate. Thus if the products are mixed to give a products stream of temperature Tav, the feedstock temperature T5, and hence temperatures T6 and T2, may be higher, provided that Tav is below TO. In Figure 3 there is shown a cylindrical catalyst bed 10 having a first end 12 and a second end 14 mounted for rotation about its longitudinal axis in the direction of the arrow. Again for simplicity, the purge stages E1 and E2 and the hydrogen flushing stage D are not shown. The bed is in the form of a honeycomb having axial through flow passages on the walls of which a catalyst material for the desired reaction is deposited. At end 12 baffles 62 sealed against the end 12 of the rotatable member 10 by means not shown, serve to separate an oxidising regeneration gas inlet region to which hot oxidising gas is supplied via line 32 from a reducing gas inlet region to which hot reducing gas is fed via line 44 and from a transfer region 60. At the second end 14 of the rotatable member 10 baffles 64 sealed against end 14 of rotatable member 10 define an oxidising regeneration gas outlet region from which spent oxidising gas is withdrawn via line 34, a reducing gas outlet region from which the spent reducing gas is withdrawn via line 46, a feedstock inlet region to which feedstock is fed via line 52, and a products outlet region from which products are withdrawn via line 54a.

That part of the rotatable member 10 between the gas inlet and outlet regions forms the oxidative regeneration zone A; that part of the rotatable member between the reducing gas inlet and outlet regions forms the reductive regeneration zone B; that portion of the rotatable member between the feedstock inlet region and the transfer region forms the first reaction zone C1; and that portion of the rotatable member between the transfer region and the products outlet region forms the second reaction zone C2.

The operation of the system is as previously described with the transfer region 60 serving to transfer mixed partially reacted feedstock from the end 12 of the first reaction zone back into end 12 of the second reaction zone, and products outlet conduit 54a serving to permit withdrawal of a mixture of products streams from various periods of the second reaction stage.

The reaction should be effected at a pressure and temperature that enables a useful conversion of the feedstock to take place. For a reaction giving an increase in the number of moles of product over the number of moles of feedstock, i.e. as in the case of propane dehydrogenation, the reaction pressure is preferably relatively low, particularly in the range 0.1 to 10 bar abs, and very conveniently at a pressure in the range 1 to 2 bar abs. Preferably the regeneration stages are effected at essentially the same pressure as the reaction stage, although it may be desirable to arrange that the hydrogen flushing stage D is maintained at a pressure lower than the reaction pressure in order to obtain greater recovery of the products stream. The catalyst material and temperatures employed will of course depend on the desired reaction. For propane dehydrogenation, the catalyst material may be a chromia/alumina mixture and may be supported on a substrate such as an alumina honeycomb. The catalyst material may also include one or more platinum group metals as a promoter. For propane dehydrogenation, the degree of conversion is limited by unfavourable equilibrium at temperatures below about 550°C and at temperatures above about 650°C by the occurrence of an undue amount of the undesired side-reactions depositing carbon. Thus the reaction temperature is preferably in the range 550-650°C. To avoid unwanted reactions in the absence of the catalyst material, the feedstock is preferably fed at a temperature below about 470°C and the cooling in the second reaction zone is preferably such that the mixture of products has a temperature below about 500°C. The invention is illustrated by the following examples.

Example 1 In this example a catalyst of chromia (8.7% by weight expressed as Cr) on a potassium modified alumina (0.7% by weight potassium expressed as K) was employed. A sample of the catalyst was charged to a reactor and reduced in hydrogen at 600°C. The reactor was then purged with an inert gas to remove the hydrogen and propane was introduced at 600°C at a pressure of 1 bar abs. The space velocity was 6000 h^~1. After various intervals the effluent gas was analysed. In the following table the conversion of propane and the selectivity to propene is quoted for various periods from the commencement of flow of propane.

It is seen to obtain a high conversion the duration of the reaction is desirably short. The yield, i.e. conversion times selectivity, decreases significantly with long reaction times.

Example 2 A catalyst of chromia (Cr₂0₃) supported on alumina and having a chromium content of 8.6% by weight was reduced in a flow of hydrogen at 600°C. After reduction, the system was purged with helium. The oxygen uptake of the catalyst was then determined by maintaining the flow of helium at 600°C at a space velocity of 3000 h^'1 and then introducing pulses of oxygen. The ratio of the number of gram atoms of oxygen adsorbed to the number of gram atoms of chromium in the system was 0.40, indicating that the Cr₂0₃ had reduced under the hydrogen flow to CrO, In another, similar, experiment a further sample of the unreduced catalyst was subjected to a flow of helium at 500°C at a space velocity of 3000 h^'1. Pulses of propane were introduced into the helium stream and the effluent analysed. Carbon dioxide and carbon monoxide were produced in such an amount that the atomic ratio of oxygen to chromium in the catalyst was decreased to 1.25. These experiments demonstrate that hydrogen and propane can decrease the oxygen to chromium atomic ratio of the catalyst by at least 0.2 .

In the following calculated examples, propane is dehydrogenated to propene at atmospheric pressure using a rotating cylindrical member 3 m diameter and 1 m long formed of an alumina honeycomb supporting a chromia as the catalyst material and having a hub of 1 m diameter. The honeycomb has axially extending flow passages of 1.4 mm hydraulic diameter giving an open area of 67%.

Example 3 In this example the catalyst material is chromia in an amount of 3.75% (expressed as Cr metal) by weight of the alumina honeycomb, oxidative regeneration is effected with air, and reductive regeneration with hydrogen. The oxidative and reductive regeneration stages A and B comprise zones representing 10% and 5% respectively of the cycle: in the oxidative regeneration stage A air at 600°C is blown at atmospheric pressure at a rate of 193 kmol/hr through the honeycomb, and in the reductive regeneration stage B, hydrogen at 300°C and atmospheric pressure is blown, in the same direction, through the honeycomb at a rate of 38 kmol/hr. In the first reaction stage C1, which represents 70% of the cycle, propane is blown at atmospheric pressure at an inlet temperature of 400°C and at a rate of 287 kmol/hr through the honeycomb in the direction counter to the air and hydrogen flows. The partially reacted gas leaving the first reaction stage C1 is passed back through the honeycomb in the second reaction stage C2 which represents 10% of the cycle. The remaining 5% of the cycle is occupied by a purge stage E1 between the second reaction stage C2 and the oxidative regeneration stage A. In this purge stage E1, process gases are purged from the honeycomb with a small flow of fuel gas. The rotational rate is such that one cycle takes 83 seconds.

The calculated yield of propene is 87 kmol/hr and the calculated temperatures (rounded to the nearest 5°C) at different proportions of the honeycomb depth (from the end at which regeneration gas enters the honeycomb) at the finish of each stage, i.e. A, B, C1 , C2, and E1 , of the cycle are set out in the following table.

The heat required to provide the heat of reaction of the propane dehydrogenation is about 10906 MJ/h. The heat provided by the oxidation of the hydrogen in the reductive regeneration stage is about 8162 MJ/h which is about 75% of the heat of reaction. The remainder of the heat required, i.e. to provide the rest of the heat of reaction and the sensible heat, results from the combustion of the deposited carbon during the oxidative regeneration stage. Example 4 In this example the catalyst material is chromia in an amount of 2.5% (expressed as Cr metal) by weight of the alumina honeycomb. Oxidative regeneration A is effected with air followed by reductive regeneration in the first reaction stage C1 by the reaction of propane with the oxidised 5 catalyst material. The air regeneration stage A represents 10% of the cycle and utilises air at 600°C and at atmospheric pressure at a rate of 128 kmol/hr. In the first reaction stage C1, which represents 75% of the cycle, propane is blown at atmospheric pressure at an inlet temperature of 400°C and at a rate of 224 kmol/hr through the honeycomb in the direction counter to the air flow. The partially reacted gas leaving the first reaction stage C1 is passed back through the honeycomb o in the second reaction stage C2 which represents 10% of the cycle. The remaining 5% of the cycle is occupied by a purge stage E1 between the second reaction stage C2 and the oxidative regeneration stage A. In this purge stage E1, process gases are purged from the honeycomb with a small flow of fuel gas. The rotational rate is such that one cycle takes 120 seconds.

The calculated yield of propene is 60 kmol/hr and the calculated temperatures (rounded to the 5 nearest 5°C) at different proportions of the honeycomb depth (from the end at which the air enters the honeycomb) at the finish of each stage, i.e. A, C1 , C2, and E1 , of the cycle are set out in the following table.

The reduction in propane yield compared with Example 3 partly results from the use of propane to complete the regeneration of the catalyst. In this example, the heat of reaction required to effect the dehydrogenation of the propane is about 7520 MJ/h. The heat produced as a result of the oxidation of the propane by the catalyst during the first part of the reaction stage and the combustion (in the subsequent oxidative regeneration stage) of the carbon formed during the oxidation of the propane amounts to about 5580 MJ/h, i.e. about 74% of the required heat of reaction. The balance of the heat of reaction plus the sensible heat is provided by the combustion, during the oxidative regeneration stage, of the carbon deposited during the propane dehydrogenation reaction.

Claims

Claims.

1. A process for performing an endothermic reaction comprising cycling a heat transfer bed, containing a solid convertible material that can be oxidised from a lower oxidation state at which the reaction occurs to a higher oxidation state, through a succession of stages: a) an oxidative regeneration stage including a step wherein an oxidising gas is passed through the heat transfer bed, whereby said convertible material is oxidised from said lower oxidation state to said higher oxidation state, and any deposited by-products of said reaction are oxidatively removed, with consequent heating of the heat transfer bed; b) a reductive regeneration stage wherein a reducing gas that is oxidised with the evolution of heat is passed through said heat transfer bed, whereby said convertible material is reduced from said higher oxidation state to said lower oxidation state with consequent further heating of the heat transfer bed; and c) a reaction stage wherein feedstock is passed through the heated heat transfer bed, whereby said feedstock is reacted to give a products stream and said heat transfer bed is cooled, each cycle of the process having a duration of less than 5 minutes.

2. A process according to claim 1 wherein said heat transfer bed comprises a rotatable member having a multiplicity of through flow passages in which, or on the wails of which, said convertible material is disposed, and said process comprises continuously rotating said rotatable member about its axis, whereby each flow passage is moved through a succession of zones including the oxidative regeneration zone, the reductive regeneration zone, the reaction zone, and then returned to the oxidative regeneration zone.

3. A process according to claim 1 or claim 2 wherein the convertible material can be reduced from the higher oxidation state to the lower oxidation state by said feedstock with the evolution of heat, and the reducing gas comprises part of the feedstock.

4. A process according to claim 3 wherein the heat transfer bed is in the form of a rotatable member having a multiplicity of through flow passages in which, or on the walls of which, said convertible material is disposed and said process comprises continuously rotating said rotatable member about its axis, whereby each flow passage is moved through a succession of zones including an oxidative regeneration zone, a reaction zone, and then returned to the oxidative regeneration zone, and the reductive regeneration stage takes place in the initial part of the reaction zone so that convertible material in the initial part of said reaction zone is reduced from the higher oxidation state to the lower oxidation state with consequent heating of that part of the heat transfer bed in that initial part of the reaction zone, and, in the remaining part of the reaction zone, said feedstock is reacted to give a products stream and that part of the heat transfer bed in said remaining part of the reaction zone is cooled.

5. A process according to any one of claims 1 to 4 wherein the reaction stage is conducted in two parts and the feedstock is passed through the heat transfer bed in the first part of the reaction stage in a first direction and in the second part of the reaction stage the partially reacted feedstock from the first part of the reaction stage is passed through the heat transfer bed in the opposite direction.

6. A process according to claim 5 wherein the oxidising regeneration gas is passed through the heat transfer bed in the direction opposite to the direction of flow of feedstock in the first part of the reaction stage.

7. A process according to claim 6 wherein a separate reductive regeneration stage is employed prior to the first part of the reaction stage and the reducing gas is passed through the heat transfer bed in the direction of flow of feedstock in the first part of the reaction stage.

8. A process according to any one of claims 1 to 7 wherein the feedstock comprises propane and the convertible material comprises a chromia containing catalyst.

9. A process according to any one of claims 1 to 8 wherein a separate reductive regeneration stage is employed prior to the reaction stage and the reducing gas is hydrogen.

10. A process according to claim 9 wherein the spent hydrogen from the reductive regeneration stage is fed through the heat transfer bed after the reaction stage as a hydrogen flushing stage, and the effluent from the hydrogen flushing stage is added to the products stream from the reaction stage.