CIRCULATING BED COFEED CONVERSION REACTOR AND METHOD FOR CONVERTING LIGHT HYDROCARBONS TO OLEFINS, GASOLINE AND METHANOL

Abstract: The invention provides a method of converting light hydrocarbons to olefins, gasoline and methanol. In the first stage, a gas containing one or more light hydrocarbons is synthesized to higher hydrocarbons including raw gasoline. However, at commercially acceptable selectivities using currently available catalyst, a majority of the light hydrocarbons remain unsynthesized. In addition, the synthesis produces significant quantities of carbon dioxide and some carbon monoxide. In this method the unsynthesized light hydrocarbons and the oxides of carbon are used to produce methanol, which can be used for several purposes including gasoline or ethylene synthesis.

Description: CIRCULATING BED COFEED CONVERSION REACTOR AND METHOD FOR CONVERTING LIGHT HYDROCARBONS TO OLEFINS, GASOLINE AND METHANOL This invention relates broadly to catalytically- enhanced conversion systems, preferably those in which conversion products are produced by contacting a feed gas with a synthesizing agent, particularly catalyst particles, at synthesizing conditions. In a preferred aspect, this invention is directed to in * situ reoxygenation of an oxidative catalyst and conversion of light hydrocarbons, e.g., methane, to higher hydrocarbons by contact with the oxidative catalyst. Further, this invention relates to the synthesis of light hydrocarbons to higher hydrocarbons, including gasoline, at conversion levels that compliment secondary product utilization. The process uses unsynthesized light hydrocarbons, together with oxides of carbon produced as secondary products, to make methanol. The methanol and/or the higher hydrocarbons may then be used to make gasoline or ethylene. Previous work on gasoline synthesis has concentrated on methane, because methane is a chemical in plentiful supply, natural gas being principally composed of methane. One known way of producing- gasoline from methane, originally developed by Mobil Oil Company, is to first reform de- sulfurized natural gas to form hydrogen, carbon monoxide and carbon dioxide, otherwise known as synthesis gas or syngas. Methanol is then synthesized from the syngas by well established low pressure techniques. Gasoline is made by passing the methanol over a suitable catalyst, as disclosed in U.S. 3,894,107, 3,931,349 and related patents. Another route, developed by the Atlantic Richfield Company (ARCO) , involves first synthesizing higher olefins from methane, using a suitable oxidizing catalyst as disclosed in U.S. Patent No. 4,777,313 and the patents listed therein. Some gasoline is produced by this synthesis, and the lower hydrocarbons produced by the synthesis can also be converted to gasoline. The main disadvantage of this method is that, at commercially acceptable selectivity, the methane conversion is relatively low. The solution adopted by ARCO was to recycle the unconverted methane, requiring a very large reactor/regenerator system. Further, the large amount of carbon dioxide produced in the synthesis was unused, being vented to the atmosphere. This means that much of the carbon from the methane is not converted to something useful. Further, the release of large amounts of carbon dioxide into the atmosphere could be environmentally detrimenta1. Generally, the catalytic conversion of low molecular weight feed gases, particularly light hydrocarbons such as methane, to higher hydrocarbons such as gasoline involves known technology. Various reactor designs have been used in the past, including fixed bed reactors and fluidized bed reactors which involve contacting the gas with an oxidative catalyst after which the catalyst is typically stripped with steam to remove residual hydrocarbons from the catalyst. The catalyst may then be reoxygenated by contacting it with air in a regenerator vessel. Following reoxygenation, catalyst is typically steam stripped once again to remove nitrogen, then recycled back to the reactor. These reactor designs have several disadvantages. For example, the separate reoxygenation vessels often add substantial costs to the system. Also, large catalyst inventories are often required. In certain systems, reactor weight hourly space velocity ("WHSV") requirements can only be met economically by operating the reactor at relatively low gas velocities, e.g., 4 feet/second or lower. Operation at reduced velocities tends to create gas backmixing, necessitating inclusion of baffles, which may result in substantial attrition of the catalyst particles, depending on the baffle design. Still other designs may have un- economically large or tall conversion reactors. The present invention addresses these collective shortcomings and offers an improved reactor design and method for converting feed gases to conversion products. This invention is broadly directed to a system for converting a reactant material, preferably a feed gas, to conversion products by contacting the reactant material with a synthesizing agent, particularly solid catalyst particles. In a preferred embodiment of the invention, the reactor can operate at significantly higher space velocity (in lbs/hr feed/lb catalyst) which tends to reduce large catalyst inventory otherwise required in conventional reactor systems. In a preferred embodiment, air or pure oxygen may be used to reoxygenate the oxidative synthesizing agent in situ while it synthesizes the conversion product; not only is less catalyst inventory required, but a separate re- oxygenator vessel is not required. In a broad aspect, the apparatus includes a reactor which comprises: (a) a first reaction zone which preferably comprises a narrowed reactor section or riser having a uniform diameter for transporting materials upward at a high velocity; (b) a second reaction zone which preferably comprises a conversion reactor having a uniform or gradually enlarging diameter, preferably cone shaped, for reducing the velocity of materials passing upward through the conversion reactor from the first reaction zone; (c) transfer means for transferring a synthesizing agent, e.g., a catalyst, for contact with a regeneration agent to form a synthesis mixture, the synthesizing agent preferably being at least partially regenerated and more preferably completely regenerated in the first reaction zone by contact with the regeneration agent, e.g., oxygen, to form the synthesis mixture; (d) propulsion means for propelling the synthesis mixture through the first reaction zone at the high first velocity; and (e) contact means for contacting the synthesis mixture with a feed gas for a sufficient time to form a conversion product, the synthesis mixture comprising the synthesizing agent and regeneration agent. In a preferred embodiment of the invention, the apparatus also comprises: (f) velocity reduction means for reducing the velocity of the synthesis mixture in the second reaction zone, the desired conversion of feed gas to conversion product being preferably completed in the second reaction zone; (g) means for accelerating the effluent stream or "product mixture" which comprises conversion product, synthesizing agent, and remaining feed gas, after the desired conversion is completed; (h) means for removing the product mixture from the second reaction zone; and (i) means for separating the synthesizing agent from the product mixture, preferably a stripper or separator. In a preferred embodiment of the invention, the system comprises a circulating bed reactor which has first and second tandemly connected reaction zones, each zone com- prising vertical reactor sections, the second section having a larger maximum diameter than the first, so that the "reactant mixture," which comprises synthesizing agent, regeneration agent, and feed gas passing through the circulating bed reactor, decelerates upon entering the second reactor section where it remains for a sufficient time to achieve desired conversion of the feed gas to conversion product. Preferably, the synthesizing agent is an oxidative catalyst which is contacted with an oxidizing agent and is reoxygenated while being transported through the two reaction zones. In the specific embodiment of the invention shown in FIG. 2 and discussed in detail below, the apparatus is specifically designed for converting feedstock comprising light hydrocarbons, e.g., methane, to higher hydrocarbons. The apparatus comprises a circulating bed reactor for reoxygenating an oxidative catalyst in situ while simul¬ taneously converting the feed gas to higher hydrocarbons. The particulate catalyst contacts the feed gas under synthesizing conditions in the reactor, preferably in the second reaction zone, to convert the methane to higher hydrocarbons. These higher hydrocarbons include, without limitation, ethane, ethylene, propylene, and aromatic hydrocarbons such as benzene and toluene. In a critical aspect, the catalyst is reoxygenated while in the reactor. Preferably, reoxygenation of the catalyst and conversion of the feed gas both occur simultaneously in the reactor as the reactant mixture proceeds rapidly upward through the two reaction zones of the reactor. Another broad aspect of the invention is directed to a method for synthesizing a feed gas comprising light hydro¬ carbons such as methane to conversion products by contacting the feed gas with solid particles at synthesizing conditions. The method comprises the steps of propelling or in some manner conveying a reactive mixture comprising solid synthesizing agent particles (preferably catalyst particles) and regeneration agent upward through a first reaction zone and toward a second reaction zone at a high first velocity, reducing the velocity of the synthesis mixture, contacting the synthesis mixture with the feed gas, and synthesizing the feed gas to conversion product at a desired conversion level in the second reaction zone. Preferably, the feed gas contacting means contacts the feed gas with the synthesis mixture to form the reactant mixture before the synthesis mixture slows down to the second velocity so that the feed gas will begin converting to conversion products before entering the second reaction zone. Preferably, the method also comprises the steps of reoxidizing the syn¬ thesizing agent in situ in the first and/or second reaction zones, removing the synthesizing agent from the second reaction zone, separating the reoxidized synthesizing agent from the conversion product, and recycling the reoxidized synthesizing agent back to the first and second reaction zones. Another aspect of this invention provides a method for converting light hydrocarbons, preferably hydrocarbons with up to four carbon atoms, to higher hydrocarbons and methanol, which comprises contacting a gas containing one or more of the light hydrocarbons at synthesizing conditions with an oxidative synthesizing agent to synthesize higher hydrocarbons from at least part of the light hydrocarbons, recovering the higher hydrocarbons, recovering carbon dioxide produced during the synthesis, mixing the recovered carbon dioxide with hydrogen and then contacting the mixture at methanol synthesizing conditions with an agent capable of synthesizing methanol, and recovering the methanol. Preferably, the gas is not recycled, i.e. is contacted one time only, the light hydrocarbons that remain un- synthesized after this single contact being recovered. The recovered light hydrocarbons, together with carbon monoxide produced in the synthesis, is then reformed to hydrogen, carbon monoxide and carbon dioxide, which are mixed with the carbon dioxide recovered after the light hydrocarbon synthesis for the purpose of synthesizing to methanol. As an alternative to recovering the light hydrocarbons, all the remaining components of the synthesizing reactor effluent after the higher hydrocarbons and water have been removed are passed through the reformer. In a preferred embodiment, the oxidative synthesizing agent comprises at least one reducible oxide of at least one metal selected from the group consisting of manganese, tin, indium, germanium, lead, antimony and bismuth, preferably manganese. It is advantageous if the oxidative synthesizing agent further comprises at least one promoter selected from the group consisting of alkali metals, alkaline earth metals and their compounds, preferably a compound of lithium. It is even more preferred that the oxidizing synthesizing agent further comprises at least one member of the group con- sisting of boron and compounds of boron. The method of this invention may particularly be used to produce gasoline from light hydrocarbons. In one embodiment, gas containing one or more light hydrocarbons is contacted at synthesizing conditions with an oxidative synthesizing agent to synthesize higher hydrocarbons including gasoline from at least part of the light hydro¬ carbons, and the gasoline and the higher hydrocarbons are separately recovered. Methanol is synthesized using the recovered carbon dioxide and reformed unsynthesized light hydrocarbons, the methanol is recovered and at least part of the methanol is converted to gasoline. Alternatively, the method may be used to produce ethylene from light hydro¬ carbons, both as one product of the light hydrocarbons synthesis and also from conversion of at least part of the methanol to ethylene. In another embodiment, the entire effluent from the synthesizing reactor is contacted at suitable conditions with an agent capable of synthesizing gasoline from olefins. This agent is preferably a suitable fluidized zeolite catalyst. After removal of water, the gasoline and other hydrocarbons higher than methane and ethane are recovered. The remaining unreacted methane, ethane and carbon dioxide may be utilized, either to make methanol or as fuel for power generation or other uses. An advantage of this method is that the carbon dioxide produced in the synthesis of higher hydrocarbons is utilized, rather than being wasted by being vented to the atmosphere with potentially adverse environmental effects. A further advantage of the preferred embodiment is that the light hydrocarbon feedstock which is not synthesized on the first pass through are not recycled to the synthesis reactor, thus permitting smaller and cheaper reactors requiring less catalyst to be used, but is still used to produce a useful product, methanol, at higher yields because of carbon dioxide utilization. The method has the further advantage of allowing flexibility in the products produced from the plant in which it is carried out, permitting a ready change in the mix of products as commercial economics change. In a general aspect, it is contemplated that the invention is applicable to any catalytically-enhanced process involving the conversion of reactant materials, particularly reactant gases, to synthesis products, parti¬ cularly those processes requiring minimal secondary reactions. Thus, although the invention is preferably directed toward the catalytically enhanced conversion of light hydrocarbons to higher hydrocarbons, it may also be used in other processes. FIG. 1 is a diagrammatic representation of the process flow, using a co-feed reactor for the light hydrocarbon synthesis. FIG. 2 is an elevation of a circulating feed and co- feed reactor design for use in the light hydrocarbon synthesis step of the process. FIG. 3 is a schematic depiction of a two stage reaction system. FIG. 4 is a diagrammatic representation of the process flow of an alternative method of utilizing the products of the light hydrocarbon synthesis. The method of this invention can be used with any source of a light hydrocarbon, preferably a hydrocarbon with up to four carbon atoms, either pure or mixed other light hydrocarbons and/or with minor amounts of inert substances. In the preferred embodiment, the feedstock is natural gas which has been field treated by desulfurizing and by removing water and a proportion of the condensable hydro¬ carbons, generally propane and higher hydrocarbons. In an alternative embodiment, refinery offgas may be used. FIG. 1 shows diagrammatically a preferred embodiment of the method of the invention. The light hydrocarbon feedstock 2 is fed to the co-feed reactor 4, where it contacts a catalyst. The catalyst is comprised of a conventional support material carrying an oxidative syn¬ thesizing agent. The composition and preparation of preferred catalysts is described in U.S. Patent Nos. 4,499,322 and 4,777,313, incorporated herein by reference, but other suitable known catalysts may be employed. Most preferably, the catalyst contains a manganese oxide as the oxidative agent, with a lithium compound and boron or a boron compound as promoters. In a cofeed reactor the catalyst is reoxygenated in situ through the introduction of a source of oxygen into the reactor, co-fed with the light hydrocarbon feedstock. In the preferred embodiment, compressed, oxygen enriched air 6 is used, but air or pure oxygen are other possible oxygen sources. Oxygen enriched air may be provided from compressed air 8 by a suitable nitrogen extraction unit 10. Any known technology for providing oxygen enriched air may be used, including the pressure swing adsorption (PSA) method of nitrogen removal and/or a cryogenic air separation plant. Preferably, the nitrogen 12 is collected and sold or otherwise utilized. Catalyst from the catalyst separation unit 12 is fed into the reactor 4, preferably in the stream of oxygen enriched air 6. The use of compressed air provides a fluidized upflow of catalyst. The light hydrocarbon feedstock 2, is introduced into the lower part of the reactor and the synthesis takes place while the catalyst and feedstock travel upward together. The preferred rate of conversion is between 25%and 35%. From the reactor 4 the effluent passes through catalyst separation 12. The separated catalyst is recycled through the reactor 4. The gases are cooled and quenched 14, and raw gasoline 15 is condensed and separated at this stage. The gases then pass to a carbon dioxide removal unit 16. After removal of the carbon dioxide the gases pass to a refrigeration unit 18 to remove unreacted light hydrocarbons and secondary synthesis products. Nitrogen may also be separated at this stage and, if so, the nitrogen 20 is either vented or collected for use or sale. The light hydrocarbons 22, together with the carbon dioxide 24 removed in stage 16, are used to synthesize methanol in a con¬ ventional methanol plant 26. The olefinic mixture from the refrigeration unit 18 is passed to a conventional olefin recovery unit 28, from which ethylene, propylene and other hydrocarbons such as LPG are recovered. FIG. 2 shows a plan elevation of the preferred design for the co-feed reactor system. The preferred embodiment uses two trains, each containing a reactor vessel 100, a reactor stripper 102 and associated cyclones. The cyclones are designed to achieve better than 99.5% recovery of 20 micron solids. While a single cyclone 104 is shown in FIG. 2, it will be understood by those skilled in the solids separation art that additional cyclones may be needed to achieve desired levels of separation of solids from effluent gases. The reactor in FIG. 2 includes a high velocity riser 113 and a lower velocity conversion reactor 100. The high velocity riser 113 should in general be large enough in diameter to provide for the transportation of a sufficient amount of synthesizing agent, e.g., catalyst particles, from riser base 112 to the conversion reactor 100 for conversion of feed gas to conversion products. The riser 113 is tandemly and preferably directly connected to the reduced velocity conversion reactor 100 which is funnel- or cone- shaped on one end, havixig a diameter which gradually expands to a maximum at 120, thereby reducing the velocity of the rapidly moving reactive mixture as it moves upward through the reactor 100. The other end of the reduced velocity conversion reactor 100 is funnel- or cone-shaped in the opposite direction, having a diameter which gradually narrows beginning at the point 120 of maximum diameter, thereby increasing the velocity of the product mixture as the mixture passes through the reactor. Although the riser 113 and conversion reactor 100 are shown in FIG. 2 as being two separate components, it is clear that a single-piece construction would work just as well. The dimensions of the reactor depend on conventional design factors and process parameters which vary from process to process and which may be supplied by persons skilled in the art to achieve the desired conversion. For example, 25%-35% conversion of methane to higher hydrocarbons should be achieved with a reactant mixture velocity in the riser of about 40 feet per second and at the widest point of the conversion reactor of about 7 feet per second. Preferably, the apparatus also comprises a high velocity exit line 122 for directing the product mixture out from the conversion reactor at an increased velocity after desired conversion of the feed gas. The velocity of the product mixture in the exit line is generally controlled by the diameter of the exit line relative to the volumetric flow rate of the mixture through the exit line 122, which in turn relates to the propulsive force exerted on the mixture through the riser 113. Thus, the smaller the diameter of the exit line 122, the greater the velocity of the mixture. The increased velocity of the mixture in the exit line tends to reduce undesirable secondary or side reactions of the feed gas with the higher hydrocarbons or other conversion products. The term "velocity" as used herein is understood to mean the average velocity of the entire particulate/gas mixture, whether it be the synthesis mixture, reactant mixture, or product mixture. The above-described apparatus shown in FIG. 2 has several advantages when used in a redox co-feed process, where conversion of a feed gas and reoxidation of the synthesizing agent take place at the same time. For example, no separate facilities are needed for reoxygenating -13- the synthesis agent. Further, the co-feed system improves space velocity requirements significantly, thus ensuring that less catalyst is needed in this apparatus than in conventional reactors. Also, because backmixing is minimized or eliminated, the tande ly connected reactor and riser sections of the invention do not require internal baffles, thus simplifying design, construction, and operation of the reactor. Generally, an advantage of this reactor design is the low temperature drop across the reactor. Thus, for conversion of methane into higher hydrocarbons, no reactor cooling is needed in spite of the exothermic nature of the reaction, and no cooling of the system is required except at high rates of conversion, e.g., greater than about 30% conversion. In operation, catalyst from the separator 102 is gravity fed through catalyst transfer line 106, which is controlled by slide valve 18. A pressure differential, which should exist between the separator 102 and the riser 113, caused by the high velocity of the synthesis mixture in the riser 113 and static head in the separator 102, tends to move the synthesizing agent through the transfer line 106 toward the riser 113 at riser base 112. Although the transfer rate of synthesizing agent may be regulated by other means for reguiaϋ-ing the flow of synthesizing agent toward the reactor, in a preferred aspect the transfer rate is regulated by slide valve 18. In riser 113, the syn¬ thesizing agent is then lifted by compressed air and/or oxygen-enriched air 110 which is injected into the riser base 112 through inlet 114 to form a synthesis mixture of air and/or oxygeh and synthesizing agent. In this propelled state, the synthesis mixture proceeds upward toward the conversion reactor at a high velocity, preferably from about 35 to about 50 feet/second, or even as high as 80 feet per second for other oxidizing catalysts besides those mentioned herein. Although the feed gas 116 may be introduced to the reactor at various points 109 along the riser 113, the feed gas 116 is preferably introduced into the reactor 100 at the point 118 at which the diameter of the reactor starts to increase. The increasing diameter of the reactor has the effect of automatically reducing the velocity of the reaction mixture, which in turn increases the catalyst residence time. The residence time needed to give the desired degree of conversion of the methane can be calculated for any set of actual conditions by those skilled in the art. In a preferred embodiment discussed above the velocity at the point of maximum diameter 120 is about 7 feet/second. Mass balance calculations for a circulating bed upflow reactor of the present invention depend on, among other things, the catalyst used, the desired conversion rate, operating temperatures and pressures, and the reactor dimensions. For example, in a hypothetical case it is contemplated that for a conversion of 35%, light hydro¬ carbons may be fed to the reactor at inlet 118 at a rate of about 1.4 lbs/hr per lb of catalyst, with a catalyst circulation rate through riser 113 and conversion reactor 100 of about 180 lbs/hr per lb/hr of methane. The catalyst bulk density is about 95 lbs/cu ft with an average catalyst particle size of about 100-150 microns. Preferably, the reactor operates at a temperature between about 1400 - 1600°F, and at a pressure between about 30 - 80 psia. At start up, and for operation at low rates of conversion, the feed gas is preheated before introduction into the reactor. However, at high rates of conversion where the reaction is exothermic , such as for methane feed, the feed gas may be at ambient temperature. After point 120 the reactor diameter decreases, thereby accelerating the effluent stream or product mixture towards the reactor outlet pipe or exit -Line 122. This acceleration helps to reduce any secondary reactions that might take place. In the preferred embodiment, the effluent stream is accelerated to about 40 feet/second. The effluent stream (or product mixture) which contains catalyst and gases including unsynthesized light hydrocarbons such as methane and synthesis products such as higher hydrocarbons, is injected tangentially into separator 102 through nozzle 124, which again accelerates the stream. In the preferred embodiment, the stream enters the separator 102 at about 55 feet/second. A bulk separation of the catalyst from the gases is achieved in the separator 102. The separator 102 of the invention broadly comprises any conventional means for separating or removing the remaining gases in the product mixture from the solid particles. Referring to Figure 2, a specific embodiment of the invention comprises a steam sparger 103 as part of the separating step. The steam sparger should be capable of stripping the catalyst particles in the separator when valve 105 is open, by removing entrained product gases from the pores of the catalyst. After the separation step, e.g., by contacting the catalyst with steam to remove the residual hydrocarbons, the catalyst exits the separator through catalyst transfer line 106 to be recycled through reactor 100. Because reoxygenation of the oxidizing agent occurs in the two reaction zones, there is no need to collect the catalyst in a dense bed for reoxygenation. The catalyst is continuously circulated through the system so that less catalyst inventory is required. After exiting the reactor, the gaseous components, which still contain catalyst fines, rise through plenum 126 and, exiting through plenum offtake pipe 128, are injected tangentially into cyclone 104. To further minimize secondary reactions, the effluent stream may be quenched by quenching means (preferably water) in the reactor outlet pipe 122 or the plenum offtake pipe 128 to about 1400°F or lower. Alternate quenching means may include cooling coils for indirect cooling. Further separation of catalyst particles may be effected in the cyclone, the separated catalyst particles being transferred to separator 102 through, for example, seal leg 130. The gas exiting from the top of the cyclone 104 may be further separated from remaining catalyst fines by being passed through a single cyclone or a series of multistage parallel cyclones to achieve the desired level of solids recovery, preferably in excess of 99.5% of solids of diameter 20 microns and larger. In a broad aspect, the apparatus of this invention comprises first and second reaction zones which are designed so that, given a constant volumetric flow rate in the upflow reactor, the high first velocity of the synthesizing agent in the first reaction zone is automatically reduced as it passes through the second reaction zone. Preferably, the volumetric flow rate is regulated in part by a variable propulsive force, such as that provided by a pressurized lift gas such as air and/or oxygen (and, if desired, feed gas) injected upward at riser base 112. The pressurized lift gas preferably also creates a low density zone in riser 113 which contributes to the circulation of catalyst through the reactor. Additionally, the flow rate may be controlled in part by introduction of feed gas at one or more locations in riser 113. Referring to the specific embodiment of the invention shown in FIG. 2, the first zone comprises riser 113 and the second zone comprises conversion reactor 100. In the preferred embodiment, the high first velocity of the synthesis or reactant mixture is automatically reduced by virtue of the gradually increasing diameter of the con¬ version reactor 100. In another broad aspect of the invention, transfer means are provided for contacting a synthesizing agent with a regeneration agent. In a preferred aspect, the syn¬ thesizing agent comprises a solid particulate catalyst and the regeneration agent comprises an oxidizing agent such as air and/or pure oxygen. In FIG. 2, the transfer means comprises a transfer line 106 which transfers the syn¬ thesizing agent from the stripper or separator to the riser/reactor where it mixes with the oxidizing agent. In still another broad aspect of the invention, propulsion means provide a high first velocity to a syn¬ thesizing agent, (e.g., a particulate catalyst). Referring to FIG. 2, this propulsion means is preferably a pressurized stream comprising the oxidizing gas directed upward through riser 113 at a high first velocity. The pressurized stream is injected at the junction of the riser and the syn¬ thesizing agent feed pipe 106 so that the synthesizing agent does not collect or form a dense bed, but instead is conveyed, preferably being immediately propelled, along with the oxidizing gas, toward the conversion zone at a high first velocity. The conversion zone begins whenever feed gas contacts the synthesizing agent, preferably at the junction of the riser 113 and reactor 100. An advantage of the reactor design of FIG. 2 is that a force for moving the catalyst through the system may be provided by the pressurized stream of gas, which may be conveniently regulated by a valve. This injection of pressurized gas, which preferably comprises an oxidizing agent to reoxygenate the oxidization catalyst, into the riser 113 also regulates both the density of the synthesis or reactant mixture in riser 113 and the pressure differential between th separator 102 and the reactor 100 to provide circulation of the catalyst through the system. In FIG. 2, the feed gas is introduced to the reactor through inlet 118, where it contacts the synthesizing agent contained in the synthesis mixture. More broadly, the apparatus may comprise any means for contacting the synthesis mixture with a feed gas, preferably just before the synthesizing agent enters the conversion reactor of the upflow reactor. However, when it is not deemed desirable to restrict the conversion of the feed gas to the reactor 100, the feed gas contacting means may be located in a different location relative to the reactor. For example, referring to FIG. 2, feed gas may be introduced proximate to the oxygen- containing gas injection inlet 114, e.g., at riser base 112, so that the synthesis of feed gas to conversion products begins in the riser 113. An important aspect of the invention is velocity reduction means for diminishing the high first velocity of the synthesizing agent to a reduced second velocity. Generally, the first velocity of the synthesizing agent may be decreased by providing a vertical upflow reactor having two reaction zones, (1) the first zone which comprises a reaction section having a first diameter and including the propulsion and regeneration means, and (2) the second zone, which comprises a conversion reactor section having a second diameter, wherein the second diameter is larger than the first diameter. The first and second reaction zones should be designed in accordance with residence time requirements which may vary depending on, among other things, the feed gas, conversion products, degree of desired conversion, and type of catalyst or other synthesizing agent. The density of the catalyst-gas mixture should be balanced between the upflow side and the downflow side. In the embodiment shown in FIG. 2, the upflow side comprises riser 113 and conversion reactor 100, while the downflow side comprises stripper/separator 102 and catalyst transfer line 106. This density balancing provides for a continuous flow of catalyst through the system with a minimum of catalyst holdup, thus reducing the amount of catalyst necessary to operate the system. The apparatus should be designed so that the velocity of the catalyst while being propelled through the riser is sufficiently high to reduce backmixing at any point. Another important aspect of the invention is the relative diameters of the first and second reaction zones which ensure that distinct high and low velocities of synthesizing agent particles are maintained. For example, in the circulating bed co-feed reactor shown in FIG. 2 , the diameter of riser 113 should be smaller than the maximum diameter of conversion reactor 100, and the velocity of the synthesis mixture in the riser 113 should be greater than its velocity in the reactor 100. In a preferred aspect of the invention, the velocity of the synthesis mixture through riser 113 is about 40 feet/second, the velocity of the reaction mixture through the main conversion reactor 100 is about 7 feet/second at point 120, and the exiting velocity is about 55 feet/second at point 122. Preferably, riser 113 is vertical, so that the* catalyst particles are evenly distributed in the synthesis mixture throughout the cross- sectional area of the riser, although it is contemplated that riser 113 may also be oriented at an angle so long as circulation of the synthesizing agent is not substantially hindered. When the apparatus is used to convert methane to higher hydrocarbons, the feed gas may be either pure or mixed methane with minor amounts o ' f other light hydrocarbons or inert substances. In a preferred embodiment, the feed gas is natural gas that has been field treated by desulfurizing and by removing water and a proportion of the condensable hydrocarbons, generally propane and higher hydrocarbons. Alternatively, refinery off gas may be used. For conversion of light hydrocarbons to higher hydro- carbons, the oxidative synthesizing agent should comprise at least one reducible oxide of at least one metal selected from the group consisting of manganese, tin, indium, germanium, lead, antimony and bismuth, preferably manganese. It is advantageous if the oxidative synthesizing agent further comprises at least one promoter selected from the group consisting of alkali metals, alkaline earth metals and their compounds, preferably a compound of lithium. It is even more preferred that the oxidizing synthesizing agent further comprises at least one member of the group con- sisting of boron and compounds of boron. Preferably, the synthesizing agent is a catalyst comprised of a conventional support material carrying an oxidative synthesizing agent. The composition and preparation of preferred catalysts is described in U.S. Patent Nos. 4,499,322 and 4,777,313, incorporated herein by reference, but other suitable known catalysts may be employed. Most preferably, the catalyst contains a manganese oxide as the oxidative agent, with a lithium compound and boron or a boron compound as promoters. To reduce or prevent catalyst deactivation, it is further preferred that all wetted surfaces in the reactor system be coated with suitable lining, such as a high alumina refractory. In the cofeed reactor shown in FIG. 2, the catalyst is reoxygenated in situ through the introduction of a source of oxygen into the reactor, co-fed with the feedstock. In a preferred embodiment, compressed oxygen-enriched air is used, but air, or steam and pure oxygen are other possible oxygen sources. Oxygen-enriched air may be provided from compressed air by a suitable nitrogen extraction unit (not shown) . Any known technology for providing oxygen enriched air may be used, including the pressure swing adsorption (PSA) method of nitrogen removal and/or a cryogenic air separation plant. Preferably, the nitrogen is collected and sold or otherwise utilized. Catalyst from a catalyst separation unit 102 may be fed into the reactor, preferably in the stream of oxygen enriched air at the junction of transfer line 106 and riser 113. The use of compressed air provides a fluidized upflow of catalyst. The light hydrocarbon feedstock is preferably introduced into the lower part of the reactor, and the synthesis takes place while the catalyst and light hydro- carbon travel upward together, in contact with one another and preferably in the second reaction zone. In accordance with this invention, the preferred rate of conversion for light hydrocarbons is between 23% and 35%. Another aspect of this invention is directed to a method for converting or synthesizing a feed gas to a conversion product by contacting the feed gas with solid particles at reaction or synthesizing conditions. Broadly, the method comprises the steps of propelling a synthesis mixture comprising solid synthesizing agent particles and regenerator agent upward from a first reaction zone toward a second reaction zone-at a first velocity, preferably in the absence of feed gas, contacting the synthesis mixture with the feed gas, reducing the velocity of the resulting reactant mixture, and synthesizing the feed gas in the reactant mixture to* conversion products in the second reaction zone. Although the solid particles in the synthesis mixture may be contacted with feed gas at any point in the first and second reaction zones, preferably such contact is made after the synthesis mixture has achieved the first velocity and before the synthesis mixture enters the second reaction zone. Preferably, the method also comprises the step of reoxidizing the synthesizing agent in situ in the first and/or second reaction zones, preferably either in the riser 113 prior to contact with the feed gas or during synthesis in the conversion reactor 100. Preferably, the method also comprises removing the product mixture from the second reaction zone, separating the reoxidized synthesizing agent from the conversion product, and recycling the reoxidized synthesizing agent back to the first and second reaction zones. In a preferred aspect, the invention is directed to an improved method for converting light hydrocarbons to higher hydrocarbons. The method comprises the steps of propelling a synthesizing agent, e.g. , a catalyst, upward through a first reaction zone, which preferably comprises a riser, to a second reaction zone, which preferably comprises a conversion reactor; and reducing the velocity of the synthesizing agent in the second reaction zone for a time sufficient to achieve desired conversion of light hydro¬ carbons to higher hydrocarbons, preferably 25% to 35%. Preferably, the method also comprises contacting the oxidative synthesizing agent with an oxidizing agent, e.g. , air, steam, or pure oxygen, and reoxygenating the synthesizing agent in situ in the first and/or second reaction zones. Preferably, the method further comprises increasing the velocity of the product mixture in the second reaction zone, removing the product mixture from the second reaction zone, stripping the synthesizing agent of higher hydrocarbons, and recycling the synthesizing agent back to the reaction zones. In a particularly preferred aspect, the invention is directed to a method of converting methane in a feed gas to higher hydrocarbons while simultaneously carrying out in situ reoxidation of the catalyst using a suitable oxidizing source in a high velocity circulating reactor. A reactor embodying aspects of this invention and shown in FIG. 2 is designed to achieve plug flow of gases and solid particles, thus avoiding the need for any internal baffling steps which are often needed in conventional reactors to prevent or reduce backmixing of reaction gases. An advantageous feature of this method includes achieving the desired conversion of methane while limiting the reaction zone residence time, thus preventing an undesirable level of secondary reactions and providing a highly selective conversion. In a further advantageous aspect of the invention, catalyst particles are circulated from a reactor zone to a separator zone in an exothermic conversion reaction, such as the conversion of methane, wherein catalyst is present in an amount sufficient to operate as a heat sink for the conversion reaction, thus providing an adiabatic operation of the reactor, and preferably eliminating any need for cooling in the reactor. The temperature rise in the reactor may be reduced by adjusting the feed gas preheat. When additional cooling is desired, the boiler feed water ("BFW") may be injected into the separator, the vaporized BFW operating to strip the catalyst of product gases. it will be apparent, of course, that many modifications may be made in the above-described embodiments without departing from the scope of the invention, which is defined by the claims below. The design of the reactor system may be slightly different than that shown in FIG. 2, yet still retain the same principles of the invention. For example, instead of a separate riser and conversion reactor, the first and second reaction zones may be incorporated into a single conversion reactor having a narrowed tubular portion of a first diameter at the bottom for propelling the synthesizing agent upward at a high velocity toward a top portion of the reactor which then widens to a second diameter to reduce the velocity of the propelled synthesizing agent in the top portion of the reactor, wherein conversion of the feed gas may take place. Also, the riser which is disposed in tandem relation to the conversion reactor in FIG. 2 may not be perfectly vertical in an alternative design, and the diameter of the conversion reactor may not be cone-shaped but may still provide the desired reduction of synthesis mixture velocity. Further, although the invention renders unnecessary an additional reoxygenating vessel for the synthesizing agent, an apparatus of the invention may comprise such a vessel as an additional feature. Instead of a cofeed reactor a two stage redox reactor may be used, in which the catalyst is reoxidized in a separate vessel from the reactor. One such reactor system is shown in FIG. 3. The light hydrocarbon feedstock 200 is fed into the lower part of reactor 202 which contains catalyst. . The effluent gases leave the reactor and pass through a series of cyclones, represented here by cyclones 204 and 206, to separate out catalyst fines. Spent catalyst leaves the reactor through tube 208 and passes through stripper 210 which strips hydrocarbons from the catalyst before it enters the regenerator 212. The spent catalyst is there contacted by a source of oxygen, which could be air or a steam/air mixture, at an elevated temperature of about 1500°F which causes the catalyst to become reoxidized. After passing through stripper 214 to remove nitrogen, the catalyst is lifted to the reactor 202 through lift feed pipe 216 by a portion of the feed stock 200 which is piped through lift feed gas supply 218. The light hydrocarbon synthesis catalyst is known to be poisoned at high temperatures by iron, therefore it is preferred that all wetted surfaces in the reactor system be coated with suitable lining, such as a high alumina refractory. After leaving the catalyst removal system, the gases are cooled to about 120°F in a series of heat exchangers. If required, final removal of solids carried over from the reactor may be effected after this cooling by contacting the gas with a water spray and then passing it through a trayed tower. In the preferred embodiment the gas stream is then compressed to about 150 psi. Much of the water and heavy hydrocarbons in the stream are liquified at this stage and removed. Large quantities of water are produced which, after suitable treatment to remove hydrocarbons and other pollutants, may be used for industrial purposes. The next step is carbon dioxide removal 16, which is preferably carried out using a conventional acid gas removal system based on aqueous methanolamine, DEA, MDEA and the like. The hydrocarbon gas after carbon dioxide removal is then cooled to about 45°F. The water and hydrocarbons condensed at these conditions are separated, and the remaining water vapor and carbon dioxide removed, preferably using a molecular sieve dehydration system. The hydrocarbon stream is then further refrigerated in a conventional cold box system 18. From this system ethylene and propylene are recovered as one product stream. This stream is further purified to remove trace compounds, compressed to about 400 psi, and then the ethylene is separated from the propylene using conventional ethylene separation techniques. In the preferred embodiment, cryogenic techniques suitable to produce polymer grade ethylene and propylene are utilized. Raw gasoline is condensed and produced as a separate stream before the cold box system. The unreacted light hydrocarbons, together with secondary synthesis products, in particular carbon monoxide, are separately recovered from the cold box system. After recovery they may still be mixed with nitrogen, but in the preferred embodiment the cold box system is designed to separate the nitrogen, which is either purged or collected. The recovered gases are compressed to about 450 psi and is then passed to the reformer to convert the light hydro¬ carbons to synthesis gas. In the preferred embodiment a conventional tubular steam reformer is used. The synthesis gas from the reformer is then cooled to about 90°F, mixed with the carbon dioxide recovered in stage 16 which has been compressed to a similar pressure, and the mixture compressed to about 1200 psi and passed to the methanol synthesis system. Any suitable known methanol synthesis system may be used, the preferred embodiment uses a Linde AG design, isothermal reactor. The gases leaving the methanol synthesis reactor are cooled to condense methanol and water, and a portion of the remaining gases are recycled to the synthesis reactor. The methanol is then separated from the water and purified by conventional techniques, preferably distillation. The degree of purification required will depend upon the eventual use for the methanol. The water may be treated as described above for further use. In an alternative embodiment shown in FIG. 4, the products of the light hydrocarbon synthesis reactor 300 are cooled by conventional cooling and quenching means 302 to about 800'F and then directed to a fluidized bed 304 containing a zeolite catalyst, preferably a ZSM-5 or equivalent catalyst. Here the olefins are converted to gasoline and LPG. The products of the reactor are cooled to separate water and all hydrocarbons higher than methane and ethane are recovered by conventional means 306. The unreacted gases, primarily methane and C0 2 , are compressed and sent to a classical steam reforming methanol unit 308. Alternatively, the unreacted gases may be directed to a gas- turbine cogeneration unit for power and/or steam production, or be used as fuel for other purposes. This embodiment has cost advantages in that much less cooling equipment is required. Further the needed degree of separation of catalyst fines is not as great, requiring less separation equipment. In either of the above-described embodiments, at least part of the methanol may be converted to gasoline, using a suitable zeolite catalyst. The preferred method is described in U.S. Patent No. 3,894,107 owned by Mobil Oil Corporation and incorporated herein by reference, and the catalysts are described in U.S. Patents Nos. 3,702,886 and 3,709,979. In the preferred embodiment a zeolite catalyst equivalent to the Mobil ZSM-5 catalyst is used, which is available commercially from a number of vendors. Alter- natively, the methanol may be used to make ethylene. This is preferably done by a two stage process. In the first stage a suitable zeolite catalyst is used to convert the methanol to DMA, then in the second stage DMA is converted to ethylene using an activated alumina catalyst. Although the primary thrust of the invention is conversion of light hydrocarbons to higher hydrocarbons, it is understood that the invention is also directed to othe processes, for example, catalytic cracking of distillate and heavier petroleum fractions to lower molecular weigh products. Also, although the preferred embodiment relate to oxidation catalyst regeneration agents, it is con¬ templated that the regeneration agent may comprise a reducing agent instead which acts upon a reactant feed by reduction rather than oxidation. In such a case, the regeneration of the reducing agent would naturally involve a different mechanism than the reoxygenation discussed herein. Additional advantages and modifications will be readily apparent to those skilled in the art. The invention in its broader aspects is therefore not limited to the specific details, representative apparatus or the illustrative example shown and described. Accordingly, departures may be made from the detail without departing from the spirit or scope of the disclosed general inventive concept.

Claims: CLAIMS : 1. A method for converting light hydrocarbons to higher hydrocarbons and methanol, which comprises: (a) contacting a gas containing one or more light hydrocarbons at synthesizing conditions with an oxidative synthesizing agent to synthesize higher hydrocarbons from at least part of the light hydrocarbons; (b) recovering the higher hydrocarbons produced during the synthesis step (a) ; (c) separately recovering carbon dioxide produced during the synthesis step (a) ; (d) mixing the recovered carbon dioxide with hydrogen or synthesis gas; (e) contacting the mixture from step (d) at methanol synthesizing conditions with an agent capable of synthesizing methanol; and (f) recovering the methanol. 2. The method of claim 1, wherein the light hydrocarbon is at least one of methane, ethane, propane, n-butane and isobutane. 3. The method of claim 1 or 2, wherein the gas is contacted once only with the oxidative synthesizing agent in step (a) . 4. The method of claim 1, 2, or 3 wherein the oxidative synthesizing agent comprises at least one reducible oxide of at least one metal selected from the group consisting of Mn, Sn, In, Ge, Pb, Sb, and Bi. 5. The method of claim 4, wherein the metal is Mn. 6. The method of claim 4, wherein the oxidative synthesizing agent further comprises at least one promoter selected from alkali metals, alkaline earth metals and compounds thereof. 7. The method of claim 6, wherein the promoter is a compound of Li. 8. The method of claim 6, wherein the oxidizing synthesizing agent further comprises boron or a compound of boron. 9. A method according to any one of the preceding claims, wherein the mixture used in step (d) is produced by: (g) recovering the light hydrocarbons that remain unsynthesized after the synthesis step (a) , and carbon monoxide produced during synthesis; (h) reforming at least part of the unsynthesized light hydrocarbons and carbon monoxide from step (g) to provide hydrogen, carbon monoxide and carbon dioxide; and (i) mixing the recovered carbon dioxide from step (c)with the products, including hydrogen, from step (h) . 10. The method according to any of the preceding claims, further comprising the step of recovering water produced during step (a) or (f) . 11. A method according to claim 9 wherein water is removed from the effluent exiting the reactor vessel used in step (a). 12. A method according to any one of the preceding claims wherein the higher hydrocarbons produced in step (a) include gasoline which is recovered therefrom and at least part of the methanol from step (f) is converted to gasoline. 13. A method according to any one of claims 1 to 11, wherein at least part of the methanol from step (f) is converted to ethylene. 14. A method of producing gasoline from light hydrocarbons, comprising: (a) passing a gas containing one or more light hydro¬ carbons through a reactor vessel at synthesizing conditions so that the gas comes into contact with an oxidative synthesizing agent to synthesize higher hydrocarbons including olefins from at least part of the light hydrocarbons; (b) contacting the effluent exiting the reactor at synthesizing conditions with an agent capable of synthesizing gasoline from olefins; (c) removing water from the product of the gasoline synthesis; (d) recovering all hydrocarbons higher than methane and ethane; (e) reforming the methane, ethane and carbon dioxide remaining after step (d) to provide hydrogen, carbon monoxide and carbon dioxide; (f) contacting the reformer output at methanol synthesizing conditions with an agent capable of synthesizing methanol; and (g) recovering the methanol. 15. The method of claim 14, wherein the agent capable of synthesizing gasoline from olefins is a suitable zeolite catalyst. 16. A method according to claim 13 or 14 comprising a feature as claimed in at least one of claims 2 to 8. 17. An upflow circulating fluidized bed reactor for converting feed gas to conversion product, comprising: (a) first and second tandemly connected reaction zones, each zone comprising vertical reactor sections; (b) means for propelling synthesizing agent upward through the first reaction zone and into the second reaction zone; (c) wherein the second vertical reactor section of said second reaction zone has a diameter which is sufficiently larger than the diameter of the first vertical reactor section of said first reaction zone to decelerate the synthesizing agent for a sufficient time to achieve a desired conversion of feed gas to - 1 -conversion product in the second reaction zone. 18. An apparatus for converting feed gas to product gas, comprising: (a) an upflow reactor comprising a first reaction zone and a second reaction zone, said first reaction zone being positioned in tandem relation with said second reaction zone, said two reaction zones being disposed to provide for the upflow of synthesizing agent from said first reaction zone to said second reaction zone; (b) transfer means for contacting a synthesizing agent with a regeneration agent, said contacting taking place in said first reaction zone of said upflow reactor; (c) propulsion means for propelling said synthesizing agent together with said regeneration agent upward from said first reaction zone toward said second reaction zone at a first velocity; (c) contact means for contacting said synthesizing agent with said feed gas; and (d) velocity reduction means for reducing the velocity of said synthesizing agent, regeneration agent, and feed gas in said second reaction zone for a time sufficient to convert the feed gas to conversion product. 19. The apparatus of claim 18 wherein said first reaction zone comprises a riser and said second reaction zone comprises a conversion reactor. 20. The apparatus of claim 19 wherein said conversion reactor has a larger diameter than the diameter of said riser. 21. The apparatus of claim 19, wherein the riser is capable of transporting a fluidized bed of solid particles from said first end to said second end of said riser. 22. The apparatus of claim 21, wherein the velocity of said transported solid particles is at least about 40 feet/sec. 23. An apparatus for converting methane to higher hydro¬ carbons by contacting the methane with synthesizing agent at synthesizing conditions and simultaneously regenerating the synthesizing agent during said methane conversion, comprising: (a) a high velocity riser having a first end and a second end, the riser being of a size sufficient to transport said synthesizing agent from said first end to said second end of said riser; (b) a low velocity conversion reactor for receiving said synthesizing agent from said high velocity riser, said conversion reactor having a first end and a second end, the conversion reactor being of a size sufficient to convert methane to higher hydrocarbons at a desired conversion level; (c) wherein the high velocity riser is disposed in tandem relation to the low velocity conversion reactor, the second end of the riser facing the first end of the conversion reactor to provide for the transport of said synthesizing agent from the high velocity riser to the low velocity conversion reactor; and (d) wherein the diameter of the low velocity conversion reactor is variable, the diameter of the conversion reactor at a point between said first end of said conversion reactor and said second end of said conversion reactor being greater than the diameter of said conversion reactor at said first end of said conversion reactor, so that synthesizing agent being transported from the high velocity riser slows down as it is transported through said conversion reactor. 24. The apparatus recited in claim 23 additionally comprising: (a) means for recycling said synthesizing agent through said high velocity riser and said low velocity conversion reactor, said recycling means including means for regulating the flow rate of synthesizing agent through said riser and reactor to control the exothermic reaction of said methane conversion; and (b) means for injecting an agent for regenerating the synthesizing agent while the synthesizing agent is providing for the conversion of methane to higher hydrocarbons, said injecting means including means for regulating the regeneration of said synthesizing agent, wherein said means for regulating the regeneration of said synthesizing agent is independently operable from said means for regulating the flow rate of said synthesizing agent. 25. The apparatus recited in claim 24 wherein said synthesizing agent comprises oxidative catalyst particles and said regeneration agent comprises an oxygen-containing gas. 26. An apparatus for synthesizing higher hydrocarbons from a methane-containing feed gas, comprising an upflow conversion reactor having a first reaction section and a second reaction section; (a) said first reaction section comprising a riser for propelling oxidative catalyst particles through said second reaction section; (b) said second reaction section comprising means for converting methane to higher hydrocarbons; (c) said first and second reaction sections comprising means for reoxidizing said oxidative catalyst particles in situ. 27. The apparatus of claim 26 additionally comprising means for reducing the velocity of said oxidative catalyst particles in said second reaction zone. 28. The apparatus of claim 27 wherein said second reaction section comprises a conversion reactor attached to said riser, said conversion reactor having a first end, a second end, and an intermediate point between said first and second ends, said first end positioned proximate said riser, and said means for reducing the velocity of said oxidative catalyst particles comprises a funnel-shaped surface in said conversion reactor, wherein the diameter of said conversion reactor increases from the first end to the intermediate point of said conversion reactor. 29. The apparatus of claim 26, additionally comprising a vessel which includes a stripper for stripping said oxidative catalyst particles before entering said first and second reaction zones. 30. The apparatus of claim 29, additionally comprising a transfer line attached to said vessel for feeding stripped catalyst particles to a riser, said riser being a part of said first reaction zone of the upflow conversion reactor, said stripped catalyst particles being regenerated in said riser. 31. The apparatus of claim 26, additionally comprising: (a) an exit line for removing said oxidative catalyst particles, unreacted methane, and higher hydro¬ carbons; (b) means for separating said catalyst from said unreacted methane and higher hydrocarbons; and (c) means for returning said separated catalyst to said first and second reaction sections. 32. A method of converting a feed gas to a conversion product by contacting the feed gas with solid particles at synthesizing conditions, comprising the steps of: (a) propelling a synthesis mixture comprising solid synthesizing agent particles and regenerator agent upward from a first reaction zone toward a second reaction zone at a first velocity; (b) contacting the synthesis mixture with a feed gas; (c) reducing the velocity of the resulting reactant mixture; and -39- (d) synthesizing the feed gas in the reactant mixture to conversion products while the reactant mixture resides in the second reaction zone. 33. The method recited j.n claim 32 wherein the synthesis mixture is contacted after the synthesis mixture has achieved the first velocity and before the synthesis mixture enters the second reaction zone. 34. The method recited in claim 32 additionally comprising the step of reoxidizing the synthesizing agent in situ in one or both of said reaction zones. 35. The method recited in claim 34, wherein the reoxidizing step is carried out in the riser prior to contacting the synthesizing agent with the feed gas; 36. The method recited in claim 35, wherein the reoxidizing step is carried out during synthesis of the feed gas to conversion product in the conversion reactor. 37. The method recited in claim 32, additionally comprising removing the product mixture from the second reaction zone, separating the reoxidized synthesizing agent from the conversion product, and recycling the reoxidized synthesizing agent back to the first and second reaction zones. 38. A method for converting light hydrocarbons to higher hydrocarbons, comprising the steps of propelling a synthesizing agent upward in a substantially vertical direction through a first reaction zone, which comprises a riser, to a second reaction zone, which comprises a conversion reactor; and reducing the velocity of the synthesizing agent in the second reaction zone for a time sufficient to achieve desired conversion of light hydro¬ carbons to higher hydrocarbons. 39. The method recited in claim 38, wherein the desired conversion is from 25% to 35% 40. The method recited in claim 38, further comprising contacting the oxidative synthesizing agent with an oxidizing agent and reoxygenating the synthesizing agent in situ in one or both of said reaction zones. 41. The method recited in claim 38, further comprising increasing the velocity of the product mixture in the second reaction zone after the desired conversion, removing the product mixture from the second reaction zone, removing higher hydrocarbons from the synthesizing agent, and recycling the synthesizing agent back to the reaction zones.

Inventors: BAINS CHARANJIT SINGH, GROSBOLL MARTIN PAUL, JAZAYERI BEHZAD, BAINS, CHARANJIT, SINGH, GROSBOLL, MARTIN, PAUL, JAZAYERI, BEHZAD
Applicants: LYONDELL PETROCHEM, LYONDELL PETROCHEMICAL COMPANY
Application year: 1991
Main classification: C10G3
IPC codes: B01J8, C07C2, C07C11, C07C29, C10L1