Title of Invention	"CONVERSION OF AN ALCOHOLIC OXYGENATE TO PROPYLENE USING MOVING BED TECHNOLOGY AND AN ETHERIFICATION STEP"
Abstract	The average propylene selectivity per on-stream cycle of an alcoholic oxygenate to propylene (OTP) process using one or more fixed beds of a dual-function oxygenate conversion catalyst is substantially enhanced by the use of a feed pretreatment step involving a catalytic etherification reaction, by switching to moving bed reactor technology in the olefin synthesis portion of the OTP flow scheme in lieu of fixed bed technology and by the selection of a catalyst on-stream cycle time of 300 hours or less. These provisions hold the build-up of coke deposits to a level which does not substantially. degrade catalyst activity, oxygenate conversion and propylene selectivity, thereby enabling maintenance of propylene average cycle yield at essentially start-of-cycle levels. The propylene average cycle yield improvement of the present invention over that achieved by the fixed bed system of the prior art is of the order of about 1.5 to 5.5 wt% or more.

Title of Invention

"CONVERSION OF AN ALCOHOLIC OXYGENATE TO PROPYLENE USING MOVING BED TECHNOLOGY AND AN ETHERIFICATION STEP"

Abstract

The average propylene selectivity per on-stream cycle of an alcoholic oxygenate to propylene (OTP) process using one or more fixed beds of a dual-function oxygenate conversion catalyst is substantially enhanced by the use of a feed pretreatment step involving a catalytic etherification reaction, by switching to moving bed reactor technology in the olefin synthesis portion of the OTP flow scheme in lieu of fixed bed technology and by the selection of a catalyst on-stream cycle time of 300 hours or less. These provisions hold the build-up of coke deposits to a level which does not substantially. degrade catalyst activity, oxygenate conversion and propylene selectivity, thereby enabling maintenance of propylene average cycle yield at essentially start-of-cycle levels. The propylene average cycle yield improvement of the present invention over that achieved by the fixed bed system of the prior art is of the order of about 1.5 to 5.5 wt% or more.

Full Text	BACKGROUND OF THE INVENTION The present invention relates generally to the use of an etherification reaction in a pretreatment step performed on the alcoholic oxygenate feed and of moving bed technology in the hydrocarbon synthesis reaction zone of an oxygenate to propylene (OTP) process. These features are coupled with a relatively short OTP catalyst on-stream cycle time to result in a level of OTP catalyst coking that does not significantly degrade the activity and propylene selectivity of the dual-function catalyst used in this OTP process. These three provisions sharply improve the average propylene yield achievable by this process over its on-stream catalyst cycle time relative to the average cycle propylene yield that is achievable by the prior art process that uses fixed bed technology for the oxygenate to hydrocarbon conversion reaction and operates with much longer catalyst cycle times. A major portion of the worldwide petrochemical industry is concerned with the production of light olefin materials and their subsequent use in the production of numerous important chemical products via polymerization, oligomerization, alkylation and the like well-known chemical reactions. Light olefins include ethylene, propylene and mixtures thereof. The art has long sought a source other than petroleum for the massive quantities of raw materials that are needed to supply the demand for these light olefin materials. Oxygenates are particularly attractive alternate source because they can be produced from such widely available materials as coal, natural gas, recycled plastics, various carbon waste streams from industry and various products and by-products from the agricultural industry. The art of making methanol and other oxygenates from these types of raw materials is well established. The prior art has revealed essentially two major techniques that are discussed for conversion of methanol to light olefins (MTO). The first of these MTO processes is represented in US-A-4,387,263. US-A-4,587,373 discloses the need to operate at a substantial pressure to make the commercial equipment of reasonable size and the diversion of a portion of the methanol feed to the dimethyl ether absorption zone to downsize the scrubbing zone. To control the amounts of undesired C4+ hydrocarbon products produced by ZSM-5 type of catalyst systems later prior art uses a non-zeolitic molecular sieve catalytic material. US-A-5,095,163; US-A-5,126,308 and US-A-5,191,141 disclose a metal aluminophosphate (ELAPO) and more specifically a silicoaluminophosphate molecular sieve (SAPO), with a strong preference for SAPO-34. The classical oxygenate to olefin (OTO) technology produces a mixture of light olefins primarily ethylene and propylene along with various higher boiling olefins. Although the classical OTO process technology possesses the capability of shifting the major olefin product recovered therefrom from ethylene to propylene by various adjustments of conditions maintained in the reaction zone, the art has long sought an OTP technology that would provide better yields of propylene relative to the classical OTO technology. Publication No. US2003/0139635A1 which was published on July 24, 2003 and describes a process for selectively converting methanol to propylene (MTP) and/or converting dimethyl ether to propylene and utilizes three reactors containing fixed beds of oxygenate conversion catalysts in a parallel flow arrangement with respect to the oxygenate feed and a serial flow arrangement with respect to the effluents of the first reactor and the second reactor. The publication teaches a dual function OTP catalyst of a pentasil-type (i.e. ZSM-5 type) having an alkali content less than 380 ppm and a zinc oxide content of less than 0.1 wt- % coupled with a restriction on cadmium oxide content of the same amount. This MTP process is further described in Rothaemel et al. "Demonstrating the New Methanol to Propylene (MTP) Process" (presented to the ERTC Petrochemical Conference in March of 2003 at Paris, France) as having an expected on-stream portion of the process cycle of 500 to 700 hours before in situ regeneration becomes necessary and shows a significant drop in conversion activity over the first five cycles. The conventional procedure for compensating for activity decay in a catalytic operation involves increasing the average reactor temperature to attempt to hold conversion in the targeted range of greater than 94% of the oxygenate charge. The problem addressed by the present invention is to enhance the average propylene selectivity of an OTP process over its on-stream cycle time and thereby diminish the requirement for recycle of olefin products other than propylene in order to compensate for lower propylene selectivity. SUMMARY OF THE INVENTION We have now discerned that propylene selectivity is a function not only of reaction conditions but also of average coke level deposited on the OTP conversion catalyst during the on-stream portion of the process cycle. Coke deposition is accelerated when the AT (i.e. T at the outlet of the reactor minus T at the inlet) across the individual OTP reaction is not carefully controlled. This is caused by the strongly exothermic nature of the desired OTP reactions which at 400 to 500° C can reach levels of 7 to 11 kcal/mole of oxygenate converted and which can cause the temperature at the tail-end of the reactor to reach levels that accelerate coke deposition and thus significantly degrade propylene selectivity. We have now found that average propylene selectivity in an OTP process can be significantly enhanced if the amount of detrimental carbonaceous deposits laid down on the catalyst during the on-stream portion of the process cycle is held to a level such that the OTP catalyst activity is not significantly diminished and oxygenate conversion is maintained over the process cycle near or at essentially start of cycle level. By the use of the term "not significantly diminished" we intend to mean that the conversion level at constant conditions is not allowed to fall more than 2 to 3% during the course of the on-stream catalyst cycle. Quite surprisingly we have also discerned that if this restriction is put on such an OTP process then the average propylene selectivity over the cycle will be improved by 1.5% to 5.5% or more. In other words if the once-through propylene selectivity of the prior art OTP process is in the range of 30% to 40% of the carbon-containing products from the OTP reactions we would anticipate the use of our invention to enable an improvement in this range to a level of at least 31.5% to 41.5% or more, thereby significantly improving the economics of the process. In sharp contrast to the OTP process of the prior art, our invention envisions shifting at least 10%, and preferably 20 to 30% or more, of the exothermic heat of reaction liberated when an alcoholic oxygenate such as methanol is converted to propylene from the OTO conversion step to a feed pretreatment step involving an etherification reaction which is designed to convert alcoholic oxygenates such as methanol to their corresponding ether (i.e. Dimethyl ether in the case of methanol) with resulting release of energy that is used to preheat and/or vaporize the feed rather than to accelerate coke deposition in the individual OTP reactors. This strategy allows tighter control of the AT across each of the OTP reactors and thus results in less detrimental coke deposition on the OTP catalyst and in lower permanent deactivation due to steam dealumination which accelerates at higher reactor temperatures. When this [Delta]T control strategy is coupled with our concept of replacing the fixed bed technology in the olefin synthesis reactors of the prior art with moving bed technology, it provides an OTP process that meets the target of not significantly disturbing initial OTP catalyst activity and propylene selectivity while operating with a catalyst circulation rate through the reactors such that the catalyst on-stream cycle time is 300 hours or less. Under these conditions the initial on- stream alcoholic oxygenate conversion of the OTP process does not drop by more than 2 to 3% thereby enabling maintenance of propylene selectivity at essentially start of run condition. The primary objective of the present invention is to provide a realistic and technically feasible solution to the problem of propylene selectivity loss during the on-stream cycle of the prior art fixed bed OTP process. A secondary objective is to improve the economics of the prior art OTP process in order to diminish the amount of olefins other than propylene that must be recycled in order to maintain overall propylene selectivity at the target level. A third objective is to control OTP reactor temperature differential by shifting at least 10% of the exothermic heat of reaction liberated when an alcoholic oxygenate such as methanol is converted to propylene from the OTP reaction step to a feed pretreatment step that catalytically etherifies feed alcoholic oxygenates to selectively produce dimethyl ether and the like ethers. Another object of the present invention is to avoid severe deactivation of the dual-function OTP catalyst utilized in the OTP process of the present invention in order to minimize the severity of the regeneration step that is needed in order to restore fresh catalyst activity, thereby minimizing hydrothermal damage and prolonging catalyst life. In one embodiment the instant invention is a continuous process for the selective conversion of an alcoholic oxygenate feed to propylene. In the first step of the process an oxygenate feed comprising alcohol contacts an acidic etherification catalyst in a first reaction zone at etherification conditions effective to form an ether-containing effluent stream and to shift at least 10% of the exothermic heat of reaction liberated when the feed is converted to propylene from the subsequent propylene synthesis step to this etherification step, thereby heating the ether-containing effluent stream to a temperature from 250 to 45O0C and producing by-product water in an amount of at least 0.5 moles per mole of alcoholic oxygenate converted. The second step adjusts the temperature of the resulting ether-containing effluent stream to a temperature from 375 to 525° C and adds sufficient diluent thereto to produce a heated mixture of ether, unreacted alcoholic oxygenate and diluent. The resulting heated mixture contacts a dual-function catalyst containing a molecular sieve known to have the ability to convert at least a portion of the oxygenates contained therein to propylene and to interconvert C2 and C4+ olefins to C3 olefins. This OTP reaction step is performed in a second reaction zone containing at least one moving bed reactor which operates at oxygenate conversion conditions effective to convert oxygenates contained in the heated mixture to propylene and at a catalyst circulation rate through the second reaction zone selected to result in an OTP catalyst on-stream cycle time of 300 hours or less. An effluent stream is then withdrawn from the second reaction zone that contains major amounts of a C3 olefin product and a water by-product and lesser amounts of a C2 olefin, C4+ olefins, C1 to C4+ saturated hydrocarbons and minor amounts of unreacted oxygenate, by-product oxygenates and aromatic hydrocarbons. In the next step this effluent stream passes to a separation zone, is cooled therein and separated into a vaporous fraction rich in C3 olefins, a water fraction containing unreacted oxygenate and by-product oxygenates and a liquid hydrocarbon fraction containing heavier olefins, heavier saturated hydrocarbons and minor amounts of aromatic hydrocarbons. At least a portion of the water fraction recovered in this separation step is then recycled to the second step to provide at least a portion of the diluent used therein. The vaporous fraction recovered in this separation step is further separated in a second separating zone into a C2 olefin-rich fraction, a C3 olefin-rich product fraction and a C4+ olefin-rich fraction. The C3 olefin-rich product fraction is then recovered as a principal product stream from the present process and at least a portion of the C2 olefin-rich fraction or of the C4+olefin-rich fraction or of a mixture of these fractions is recycled to the propylene synthesis step in order to interconvert these materials into additional quantities of the desired propylene product. In another step of the process a stream of coke-containing dual-function catalyst particles is withdrawn from the second reaction zone, oxidatively regenerated with an oxygen-containing stream in a regeneration zone and a stream of regenerated catalyst particles is returned to the second reaction zone in order to provide regenerated catalyst for circulation therethrough. A second embodiment involves a continuous process for the selective conversion of an alcoholic oxygenate feed to propylene as described in the first embodiment wherein the dual-function catalyst contains a zeolitic molecular sieve having a structure corresponding to ZSM-5, or an ELAPO molecular sieve having a structure corresponding to SAPO-34 or a mixture of these materials. Another embodiment comprises a continuous process for selective conversion of an alcoholic oxygenate feed to propylene as described above in the first embodiment wherein the second reaction zone contains at least 3 moving bed reactors which are connected in a serial flow or parallel flow configuration with respect to heated mixture fed thereto and in a serial flow configuration with respect to the stream of catalyst particles that passes therethrough. A highly preferred embodiment of the present invention comprises a continuous process for the selective conversion of an alcoholic oxygenate feed to propylene as described above in the first embodiment wherein the alcoholic oxygenate feed essentially comprises methanol. A high propylene yield embodiment of the instant process involves the continuous process for selective conversion of an alcoholic oxygenate feed to propylene as described in any of the previous embodiments wherein the liquid hydrocarbon fraction recovered in the first separation step is further separated into a C4 to C6 olefin-rich fraction and a naphtha product fraction and at least a portion of the C4 to C6 olefin-rich fraction is recycled to the OTP conversion step in order to interconvert these heavier olefins into propylene. BRIEF DESCRIPTION OF THE DRAWING The FIGURE is a process flow diagram of a preferred embodiment of the present invention. TERMS AND CONDITIONS DEFINITIONS The following terms and conditions are used in the present specification with the following meanings: (1) A "portion" of a stream means either an aliquot part that has the same composition as the whole stream or a part that is obtained by eliminating one or more readily separable component therefrom. (2) An "overhead" stream means the net overhead recovered from the specified zone after recycle of any portion to the zone for reflux or any other reason. (3) A "bottom" stream means the net bottom stream from the specified zone obtained after recycle of any portion for purposes of reheating and/or reboiling and/or after any phase separation. (4) The term "light olefins" means ethylene, propylene and mixtures thereof. (5) The expression "OTP" process means a process for converting an alcoholic oxygenate to propylene and in a preferred embodiment when the alcoholic oxygenate is methanol the OTP process is referred to as an MTP process herein. (6) The term "catalyst on- stream cycle time" means the length of time the catalyst particle is exposed to feed at conversion conditions before withdrawal from the reaction zone for regeneration. (7) The term "propylene average cycle yield" means the total propylene yield during the catalyst on- stream cycle time divided by the total amount of oxygenate feed converted during the catalyst on-stream cycle time. (8) The term "dual-functional" means that the OTP catalyst catalyzes both the OTP reactions and the olefin interconversion reactions necessary to convert C2 and C4+ olefins to propylene. STATEMENT OF THE INVENTION The present invention relates to a continuous process for selective conversion of an alcoholic oxygenate feed to propylene comprising the steps of: a) contacting the feed with an acidic etherification catalyst in a first reaction zone at etherification conditions effective to form an ether-containing effluent stream and to shift at least 10% of the exothermic heat of reaction liberated when the feed is converted to propylene from the subsequent propylene synthesis step to this etherification step, thereby heating the ether-containing effluent stream to a temperature from 250 to 450° C and producing by-product water in an amount of at least 0.5 mole per mole of alcoholic oxygenate converted; b) adjusting the temperature of the resulting ether-containing effluent stream to a range of 375 to 525° C and adding diluent thereto to produce a heated mixture of ether, unreacted alcoholic oxygenate and diluent; c) reacting the resulting heated mixture with dual-function catalyst particles containing a molecular sieve, having the ability to convert the oxygenates contained therein to C3 olefin and to inter convert C2 and C4+ olefins to C3 olefin, in a second reaction zone containing at least one moving bed reactor wherein the reaction zone is operated at oxygenate conversion conditions effective to convert the oxygenates contained in the mixture to propylene and at a catalyst circulation rate through the second reaction zone selected to result in a catalyst on- stream cycle time of 300 hours or less to produce a propylene-containing effluent stream containing major amounts of a C3 olefin product and water, lesser amounts of a C2 olefin, C4+ olefins and C1 to C4+saturated hydrocarbons and minor amounts of unreacted oxygenate, by product oxygenates and aromatic hydrocarbons; d) passing the propylene-containing effluent stream to a separation zone and therein cooling and separating this effluent stream into a vaporous fraction rich in C3 olefin, a water fraction containing unreacted oxygenates and by-product oxygenates and a liquid hydrocarbon fraction containing heavier olefins, heavier saturated hydrocarbons and minor amounts of aromatic hydrocarbons; e) recycling a portion of the water fraction recovered in step d) to step b) to provide a portion of the diluent added therein; f) separating the vaporous fraction into a C2 olefin-rich fraction, a C3 olefin-rich product fraction and a C4+ olefin-rich fraction; g) recycling a portion of the C2 olefin-rich fraction or of the C44 olefin-rich fraction or of a mixture of these fractions to step c); and h) withdrawing coke-containing dual-function catalyst particles from the second reaction zone, oxidatively regenerating the withdrawn catalyst particles in a regeneration zone and returning a stream of regenerated catalyst particles to the second reaction zone. DETAILED DESCRIPTION OF THE INVENTION The feedstream charged to the instant OTP process comprises one or more alcoholic oxygenates in admixture with an aqueous diluent. The concentration of the one or more alcoholic oxygenates in the feedstream is preferably at least 70 wt-% while a concentration of 90%+ is preferred with best results obtained when the alcoholic oxygenate in the feedstream is at a level of 95 wt-% or more. The alcoholic oxygenate that is present in the feedstream preferably contains at least one oxygen atom and about 1 to 6 carbon atoms with best results usually obtained with an aliphatic alcoholic oxygenate containing 1 to 4 carbon atoms. Suitable alcoholic oxygenates for use in the instant process include saturated straight or branch chain aliphatic alcohols and their unsaturated counterparts. Representative of suitable aliphatic alcohols include methanol, ethanol, isopropyl alcohol, normal propyl alcohol, butyl alcohol, isobutyl alcohol, t-butyl alcohol, amyl alcohol and mixtures of these aliphatic alcohols. Methanol is the most preferred alcoholic oxygenate for use in the feed and particularly preferred results are obtained with a feedstream which contains a relatively high concentration of methanol. Commercial feedstreams of a so-called "crude" methanol stream usually contains 70 to 80 wt-% methanol in admixture with water and of a purified methanol stream typically contains methanol in a concentration of 95 wt-% or more in admixture with trace amounts of water. In the first reaction step of the instant process the feedstream is contacted with an acidic etherification catalyst in a first reaction zone at etherification conditions effective to form an effluent stream containing at least one ether corresponding to the alcoholic oxygenate present in the feedstream. If the feedstream contains only methanol the etherification forms dimethyl ether and if the feedstream contains only ethanol diethylether (DEE) if formed with corresponding simple and mixed ethers formed in the case where the feedstream contains a mixture of alcoholic oxygenates. A significant feature of the invention is performing the etherification step at a conversion level sufficient to shift at least 10%, and preferably 20 to 30% or more, of the exothermic heat of reaction liberated when the feed is converted to propylene from the subsequent propylene synthesis step to this etherification step. From a heat integration standpoint it is highly preferred to operate the etherification reaction to preheat and/or vaporize at least a portion of the feedstream and to cut the exothermicity associated with the downstream processing of the resulting ether-rich stream in the propylene synthesis (OTP) reaction zone. The reaction of an ether such as Dimethyl ether to produce propylene and other olefins is significantly less exothermic than the direct conversion of an alcohol to an olefin in the OTP reaction zone. Accordingly the feed pretreatment step of etherification captures a significant portion of the exothermic heat released in the overall conversion of an alcoholic oxygenate to the desired olefinic hydrocarbons, thereby providing a convenient source of heat for bringing the feedstream up to the desired reaction for the propylene synthesis step and/or vaporizing a significant portion of the feedstream to facilitate a vapor phase reaction in the olefin synthesis step. An unobvious benefit of this shifting of a significant portion of the heat released during the highly exothermic conversion of an alcohol to olefinic hydrocarbons is the enabling of pretreatment step to provide sharper control over the AT experienced in the olefin synthesis reaction step. Since the coke forming side reaction is accelerated at higher temperatures and the catalyst dealumination reaction is similarly accelerated the ability to control the temperature increase in the OTP reaction zone suppresses not only this coke formation side reaction that degrades high propylene selectivity but also this catalyst dealumination reaction. A significant benefit of the instant invention is thus that the selectivity of the OTP reaction to the desired propylene olefin is enhanced by the fact that tight control can be maintained over the average temperature increase experienced in the olefin synthesis reaction zones. The feed pretreatment step or etherification step of the present invention is catalytic and requires the presence of mildly acidic heterogeneous catalyst in order to accelerate the reaction. Suitable catalysts for the etherification step of the present invention are mildly acidic porous solids having a surface area of about 10 to 500 m2/g and include the following materials: (1) silica or silica gel, clays and silicates including those synthetically prepared and naturally occurring, which may or may not be acid treated, for example, attapulgus clay, china clay, diatomaceous earth, fuller's earth, kaolin, bentonite, kieselgurh, etc. (2) refractory inorganic oxides such as alumina, titanium dioxide, zirconium dioxide, chromium oxide, beryllium oxide, vanadium oxide, cesium oxide, hafnium oxide, zinc oxide, magnesia, boria, thoria, silica-alumina, silica-magnesia, chromia-alumina, alumina-boria, silica-zirconia, etc. (3) crystalline zeolitic aluminosilicates such as naturally occurring or synthetically prepared mordenite and/or faujasite, ZSM-5, ZSM-11, or ZSM-12, either in the hydrogen form or in a form which has been treated with multivalent cations; and (4) combinations of materials from one or more of these groups. The preferred porous catalysts for use in the present invention are refractory inorganic oxides, with best results obtained with an alumina material. Suitable alumina materials are the crystalline aluminas known as gamma-, eta- and theta-alumina, with gamma- or eta-alumina giving best results. In addition, in some embodiments, the alumina carrier material may contain minor proportions of other well known refractory inorganic oxides such as silica, zirconia, magnesia, etc.; however, the preferred catalyst is substantially pure gamma- or eta-alumina. Preferred catalysts have an apparent bulk density of about 0.3 to about 0.9 g/cc and surface area characteristics such that the average pore diameter is about 20 to 300 Angstroms, the pore volume is about 0.1 to about 1 cc/g and the surface area is about 100 to about 500 m2/g. In general, best results are typically obtained with a gamma-alumina carrier material which is used in the form of spherical particles having: a relatively small diameter (i.e. typically about 1/16 inch), an apparent bulk density of about 0.3 to about 0.8 g/cc, a pore volume of about 0.4 ml/g, and a surface area of about 150 to about 250 m2/g. The etherification conversion conditions utilized in the present invention include a temperature from 200 to 375°C with a preferred range of 275 to 35O0C, an inlet pressure from 136 to 1136 kPa (5 to 150 psig) with best results achieved at a pressure from 136 to 791 kPa (5 to 100 psig). For the lower boiling alcoholic oxygenates such as methanol and ethanol the temperature and pressure used in the etherification step are usually chosen to maintain substantial vapor phase conditions or a mixed liquid/vapor condition. The weight hourly space velocity (WHSV) which is utilized in the etherification step is calculated on the basis of the mass hourly flow rate of the sum of the mass of the alcoholic oxygenate material contained in the feedstream plus any alcoholic oxygenate material contained in recycle streams divided by the mass of the etherification catalyst present in the etherification conversion zone. Ordinarily good results are obtained with WHSV of about 0.1 to 10 hrs-1 with best results ordinarily achieved in the range of about 0.5 to 5 hrs-1. In accordance with the present invention these conversion conditions are adjusted to achieve a release of at least 10%, and preferably about 20 to 30% or more, of the exothermic heat of reaction liberated when the alcoholic oxygenate contained in the feed is converted to propylene and other olefins as previously explained. As a result of the exothermic heat liberated in the etherification step the effluent stream withdrawn therefrom is recovered at a temperature from 250 to 45O0C and contains by-product water in amounts of at least 0.5 moles per mole of alcoholic oxygenate converted in the etherification step. This effluent stream from the etherification step also contains substantial amounts of one or more ethers corresponding to the one or more alcoholic oxygenates present in the feedstream. For example, with preferred alcoholic oxygenate of methanol the effluent stream from this step will contain substantial quantities of dimethyl ether. In the next step of the present invention the temperature of the resulting ether- containing effluent stream is adjusted to a range from 375 to 525° C either by adding or subtracting the necessary thermal energy from this effluent stream utilizing suitable heating and/or cooling means well known to those skilled in the art. It is also a preferred practice during this effluent temperature adjustment step to add sufficient diluent to this effluent stream to result in a mole ratio of diluent to oxygenate equivalents (i.e. the unreacted alcoholic oxygenates plus the ether content of the effluent stream) of about 0.5:1 to 5:1 with best results usually achieved at a diluent ratio of about 1:5 to 2:1. It is of course within the scope of the present invention to perform this diluent adjustment step at least in part after the ether-containing effluent stream has been passed through one or more of the downstream propylene synthesis reaction zones in the case where a multibed propylene synthesis scheme like that shown in the attached drawing is employed. After completing adjustments to the temperature and diluent content of the effluent stream withdrawn from the etherification step the resulting heated mixture of ether, unreacted alcoholic oxygenate and diluent passes to the propylene synthesis reaction step of the present invention (the OTP conversion step) to selectively prepare the desired propylene product. In the OTP conversion step of the present invention, the resulting heated mixture is catalytically converted to hydrocarbons containing aliphatic moieties such as - but not limited to - methane, ethane, ethylene, propane, propylene, butylene, and limited amounts of other higher aliphatics by contacting the heated mixture with a dual-function OTP catalyst. Presence of a diluent is useful to maintain the selectivity of the OTP catalyst to produce light olefins, particularly propylene. The use of a diluent such as steam can provide certain equipment cost and thermal efficiency advantages as well as lowering the partial pressure of the oxygenate reactants, thereby increasing selectivity to olefins. The phase change between steam and liquid water can also advantageously transfer heat between the feed to the OTP reaction zone and the OTP reactor effluent, and the separation of the steam diluent from the product requires only a simple condensation step to separate water from the hydrocarbon products. [0024] As previously stated, the amount of diluent used in the OTP reaction step will preferably the range from about 0.5:1 to 5:1 moles of diluent per mole of oxygenate equivalent and preferably 0.5:1 to 2:1 in order to lower the partial pressure of the oxygenates to a level which favors production of propylene. Some of the embodiments of the present invention envision recycling of certain olefinic streams that contain significant amounts of olefins other than propylene and saturated hydrocarbons. These recycle streams will furnish diluent to the OTP reaction zone and may reduce the amount of diluent that must be added to the OTP reaction zone to achieve the target diluent to oxygenate mole ratio once the OTP reaction zone is started. In the most preferred case water is used as the primary diluent and the amount of water charged to the OTP reaction zone during startup will diminish in proportion to the amount of other diluents that are recycled to this reaction zone. The conversion conditions used in the OTP reaction zone are carefully chosen to favor the production of propylene from the oxygenates. An oxygenate conversion temperature range of 350° to 600° C is effective for the conversion of oxygenate over the known oxygenate conversion catalysts. The lower temperature in this range favor the production of propylene and upper temperature of the range favor the production of ethylene. Preferred inlet temperatures to the OTP reaction zone range from 375° to 525° C and more preferably from 400° C to 500° C. The temperature increase across each of the OTP reactors is preferably held to a range of 10° to 60° C to minimize hydrothermal deactivation and to avoid the acceleration of coke deposition on the catalyst that occurs when the tail end temperature individual reactors builds to levels beyond those contemplated by the present invention. Most methods for controlling the temperature increase in a reaction zone involve utilization of multiple beds of catalyst in separate reactors with inner-bed or inter-bed cooling utilizing appropriate heat exchange and/or addition of relatively cool recycle streams, or portions of the oxygenate feed and/or the diluents that is utilized in the zone. The present invention contemplates the use of lighter and heavier olefin interconversion reactions which are mildly endothermic to help control reactor temperature increases to the specified range. The preferred mode of operation of the instant invention uses at least 3 moving bed reactors with interbed quench achieved at least in part by the use of relatively cool recycle streams to provide additional quantities of reactants and diluent. The oxygenate to propylene conversion step is effectively carried out over a wide range of pressures including inlet total pressures of about 0.1 to 100 atm (10.1 kPa to 10.1 mPa) and more typically between about 136 to 1136 kPa (5 to 150 psig). Since the formation of lighter olefins like propylene are favored at low pressure conditions the preferred inlet pressure ranges from 136 to 343 kPa (5 to 35 psig). The contact time of the reactants with the dual-function catalyst is ordinarily measured in relative terms of a Weight Hourly Space Velocity (WHSV) calculated on the basis of a mass hourly flow rate of the sum of the mass of oxygenate reactants passed to the OTP conversion zone plus the mass of any reactive hydrocarbon material present in the f eedstream or any of the recycle streams divided by the mass of the dual-function catalyst present in the OTP conversion zone. Weight Hourly Space Velocity (WHSV) in the OTP conversion zone ranges from about 0.1 to 100 hr-1, with a preferred range of 0.5 to 20 hr-1, and with best results ordinarily attained in the range of 0.5 to 10 hr-1. In the oxygenate to propylene conversion step it is preferred to use a dual-function catalyst having the capability of converting oxygenates to propylene as well as the capability of interconverting olefins other than propylene to propylene. Any known catalytic materials known having the capability to catalyze these two reactions are suitable catalysts for the present invention. The preferred dual-function catalyst contains a molecular sieve as the active ingredient and more specifically the molecular sieve has relatively small pores. The preferred small pore catalysts are defined as having pores at least a portion, preferably a major portion, of which have an average effective diameter characterized such that the adsorption capacity (as measured by the standard McBain-Bakr gravimetric adsorption method using given adsorbate molecules) shows good adsorption of oxygen (average kinetic diameter of about 0.346 run) and negligible adsorption of isobutane (average kinetic diameter of about 0.5 nm). More preferably the average effective diameter is characterized by good adsorption of xenon (average kinetic diameter of about 0.4 nm) and negligible adsorption of isobutane, and most preferably, by good adsorption of n-hexane (average kinetic diameter of about 0.43 nm) and negligible adsorption of isobutane. Negligible adsorption of a given adsorbate is adsorption of less than three percent by weight of the catalyst and good adsorption of the adsorbate is over three percent by weight of the adsorbate based on the weight of the catalyst. Certain of the catalysts useful in the present invention have pores with an average effective diameter of less than 5 A. The average effective diameter of the pores of preferred catalysts is determined by measurements described in D. W. Breck, ZEOLITE MOLECULAR SIEVES by John Wiley & Sons, New York (1974). The term "effective diameter" denotes that occasionally the pores are irregularly shaped, e.g., elliptical, and thus the pore dimensions are characterized by the molecules that can be adsorbed rather than the actual dimensions. Preferably, the small pore catalysts have a substantially uniform pore structure. Suitable dual-function catalysts may be chosen from among zeolitic molecular sieves and non-zeolitic molecular sieves. Zeolitic molecular sieves in the calcined form may be represented by the general formula: (Formula Removed) where Me is a cation, x has a value from about 2 to infinity, n is the cation valence and y has a value of about 2 to 100 or more and more typically about 2 to 25. Zeolites which may be used include chabazite - also referred to as Zeolite D, clinoptilolite, erionite, ferrierite, mordenite, Zeolite A, Zeolite P, ZSM-5, ZSM-11, and MCM- 22. Zeolites having a high silica content, (i.e., Those having silica to alumina ratios greater than 10 and typically greater than 100 with best results achieved at a silica to alumina mole ratio of about 250:1 to 1000:1 are especially preferred.) One such high silica zeolite having the structure of ZSM-5 is silicalite, as the term used herein includes both the silicapolymorph disclosed in US-A-4,061,724 and also the F-silicate disclosed in US-A-4,073,865. Non-zeolitic molecular sieves include molecular sieves which have the proper effective pore size and are embraced by an empirical chemical composition, on an anhydrous basis, expressed by the empirical formula: (Formula Removed) where EL is an element selected from the group consisting of silicon, magnesium, zinc, iron, cobalt, nickel, manganese, chromium and mixtures thereof, x is the mole fraction of EL and is at least 0.005, y is the mole fraction of aluminum and is at least 0.01, z is the mole fraction of phosphorous and is at least 0.01 and x + y + z = 1. When EL is a mixture of metals, x represents the total amount of the element mixture present. Preferred elements (EL) are silicon, magnesium and cobalt with silicon being especially preferred. The preparation of various ELAPOs may be found in US-A-5,191,141 (ELAPO); US-A-4,554,143 (FeAPO); US-A-4,440,871 (SAPO); US-A-4,853,197 (MAP05 MnAPO, ZnAPO, CoAPO); US-A-4,793,984 (CAPO); US-A-4,752,651 and US-A-4,310,440. A preferred embodiment of the invention is one in which the element (EL) content varies from about 0.005 to about 0.05 mole fraction. If EL is more than one element, then the total concentration of all the elements is between about 0.005 and 0.05 mole fraction. An especially preferred embodiment is one in which EL is silicon (usually referred to as SAPO). The SAPOs which can be used in the instant invention are any of those described in US-A- 4,440,871; US-A-5,126,308, and US-A-5,191,141. SAPO-34 is preferred. [0034] The preferred OTP conversion catalyst is preferably incorporated into porous solid particles in which the catalyst is present in an amount effective to promote the desired OTP reactions, hi one aspect, the porous solid particles comprise a catalytically effective amount of the molecular sieve catalyst and at least one matrix material, preferably selected from the group consisting of binder materials, filler materials, and mixtures thereof to provide a desired property or properties, e.g., desired catalyst dilution, mechanical strength, and the like to the solid particles. Such matrix materials are porous in nature and may or may not be effective to help promote the desired OTP conversion. Filler and binder materials include, for example, synthetic and naturally occurring substances such as metal oxides, clays, silicas, aluminas, silica-aluminas, silica-magnesias, silica-zirconias, silica-thorias, silica-berylias, silica-titanias, silica-alumina- thorias, silica-alumina-zirconias, aluminophosphates, mixtures of these and the like. If matrix materials, e.g., binder and/or filler materials, are included in the catalyst composition, the non-zeolitic and/or zeolitic molecular sieve catalyst preferably comprise about 1 % to 99%, more preferably about 5% to about 90% and still more preferably about 5% to about 50%, by weight of the total composition. The most preferred zeolitic dual-function catalyst for use in the OTP reaction step of the present invention is a zeolite having the structural configuration of ZSM-5, sometimes in the literature referred to as having a "pentasil-type" structure and disclosed in US 2003/0139635Al. A borosilicate zeolite having the ZSM-5 structural configuration is disclosed as a particularly preferred dual-function catalyst in US-A-4,433,188. The use of a zeolitic catalyst having the mordenite structural configuration is disclosed in GB-A-2171718. Another particularly preferred class of dual-function catalyst for use in the present invention is the ELAPO molecular sieves. These materials catalyze both the direct conversion of oxygenates to light olefins and the interconversion of olefins to a desired product olefin as taught in US-A-4,677,243 and US-A-4,527,001. US-B-6,455,749 teaches the use of silicoalumino phosphate catalyst (SAPO) as a dual-function catalyst and specifically has a preference for SAPO-34 and teaches both the use of a SAPO-34 type of catalyst system and a ZSM-5 bound with silicalite type of catalyst system for use in the interconversion of C4 olefins to other olefins. Best results with a non-zeolitic catalytic system are obtained when SAPO-34 is utilized as the dual-function catalyst. Best results with a zeolitic material are obtained with a ZSM-5 type of material. A particularly preferred feature of the invention uses a mixture of a zeolitic catalyst system with a non-zeolitic catalyst system. This mixed catalyst embodiment can use either a physical mixture of particles containing the zeolitic material with particles containing the non-zeolitic material or the catalyst or a formulation that mixes the two types of material into a binder in order to form particles having both ingredients present therein. In either case the preferred combination is a mixture of ZSM-5 with SAPO-34 in relative amounts such that SAPO-34 comprises 30 to 70 wt-% of the molecular sieve portion of the mixture with a value of about 45 to 55 wt-% being especially preferred. The present invention uses moving bed technology in the OTP conversion step to enhance the selectivity of the process for propylene production. The use of moving bed technology is shown in US-A-5,157,181. Suitable moving bed reactors and regeneration systems have been widely employed commercially for example in the catalytic reforming of naphtha fractions to increase octane number and to facilitate the dehydrogenation of light paraffins to make olefins. Moving bed reaction zones can have various configurations. The catalyst particles can be introduced to an upper section of the reaction zone and fed by gravity through the entire volume of the reaction zone, wherein the catalyst is contacted with the feedstream either in a countercurrent direction to the catalyst movement or in a concurrent direction. In a preferred aspect of the present invention the feedstream flow is countercurrent to the catalyst flow with the feedstream introduced into a lower portion of the reaction zone and withdrawn from an upper portion thereof. This preferred configuration can provide substantial advantages in OTP conversion reactions by contacting the feedstream with partially deactivated catalyst during the initial stages of the conversion when the driving force is high and more active catalysts during the subsequent stages of the conversion when the driving force is lower. More typically the catalyst particles are introduced into an annulus defined by concentric catalyst retaining screens that runs through the reaction zone, wherein the catalyst particles travel down through the annulus and are withdrawn from a lower section of the OTP reaction zone. The feedstream is introduced either to the upper or lower section of the reaction zone and is passed across the annulus generally in a direction transverse to the catalyst flow. The radial bed configuration can provide low pressure drop across the reaction zone, hence good flow distribution. During the OTP conversion zone traversal, a carbonaceous material, i.e., coke, is deposited on the catalyst as it moved downward through the reaction zone. The carbonaceous deposit material reduces the number of active sites on the catalyst thereby affecting the extent of the conversion and the selectivity to propylene. Thus during the moving bed conversion process a portion of the coked catalyst is withdrawn from the OTP reaction zone and regenerated in a regeneration zone to remove at least a portion of the carbonaceous material. Preferably the carbonaceous material is removed from the catalyst by oxidative regeneration wherein the catalyst which is withdrawn from the reactor is contacted with an oxygen-containing gas at sufficient temperature and oxygen concentration to allow the desired amount of the carbonaceous materials to be removed from the catalyst. Depending upon the particular catalyst and conversion it can be desirable to substantially remove the carbonaceous material, e.g., to less than 1 wt-% and more preferably less than 0.5 wt-%. In some cases it is advantageous to only partially regenerate the catalyst, e.g., to remove from about 30 to 80 wt-% of the carbonaceous material. Preferably, the regenerated catalyst will contain about 0 to 20 wt-% and more preferably from about 0 to 10 wt-% carbonaceous material. Preferably in most instances when relatively large concentrations of carbonaceous material (i.e. coke) are present on the catalyst, that is, greater than about 1 wt-% carbonaceous material on the catalyst, carbon burn-off occurs with an oxygen-containing gas stream containing a relatively low concentration of oxygen. Preferably the oxygen content of oxygen containing gas is regulated through the use of inert gases or the recycling of flue gas materials to maintain an initial oxygen level of from about 0.5 to 2 vol.-%. By using the small concentration of oxygen, it is possible to reasonably control the oxidation of carbonaceous materials upon the catalyst without allowing excessively high temperatures to occur, thereby preventing the possibility of permanently hydrothermally damaging the molecular sieve catalyst. Temperatures used during regeneration should be in the range from 400° to 700° C with best results obtained at about 500°to 650° C. The regeneration zone is preferably configured as a moving bed zone similar to the moving bed configuration used in the OTP reaction zone with coked catalyst fed to an upper portion of the regeneration zone and passed by gravity feed through the regeneration zone wherein the carbonaceous material is removed and the regenerated catalyst is withdrawn from a lower section of the regeneration zone and recirculated to the OTP reaction zone. DETAILED DESCRIPTION OF THE DRAWING The attached drawing shows a schematic representation of a flow scheme of the present invention in which a first reaction zone, 30, contains a fixed bed of etherification catalyst and a second reaction zone comprising reactors 1, 2 and 3 contains dual-function OTP conversion catalyst in a moving bed configuration. The three moving bed reactors 1, 2 and 3, are in serial configuration both with respect to the oxygenate feed and to the flow of catalyst through the reactors. The illustrated vertical cross section of the reactors shows catalyst descending through an annulus 33 which is maintained with appropriate concentric catalyst retaining screens. All three moving bed reactors operate with the charge to the reactor flowing in countercurrent relationship to the descending stream of catalyst particles. Preferably all three moving bed reactors have the oxygenate-containing feed stream flowing from the outside of the catalyst annulus 33 into the center portion from which the effluent stream is recovered. The drawing shows the flow of feed materials, intermediate materials and product materials by solid lines and the flow of catalyst to and from the reaction zones by dotted lines. A transport medium, preferably steam, nitrogen or any other available inert diluents transports the catalyst. The preferred catalyst transport medium is steam due to its substantial presence in the OTP reaction zone. The etherification catalyst used in the first reaction zone, zone 30, is an acidic porous alumina material used in the form of spherical particles having a diameter of about 1.6 mm (0.0625 in) comprising gamma-alumina and having a surface area of about 180 m2 g. The catalyst utilized in the second reaction zone comprising reactors 1, 2 and 3 is selected from the dual-function catalyst previously described and is utilized in a spherical form having a effective diameter of about .5 to 5 mm with a diameter of 1.6 mm being especially preferred. The total amount of OTP catalyst is preferably divided equally among the three reactors. The number of reactors in the second reaction zone is selected to hold the conversion conditions in the individual reactors at conditions which enhance the yield of propylene by limiting the temperature differential across the individual reactors to 60° C or less thereby avoiding a shift in the yield structure towards ethylene and simultaneously minimizing the coke formation on the catalyst which accelerates rapidly as previously explained. Line 8 charges a methanol-rich stream containing 95 wt-% methanol in admixture with 5 wt-% water to the process. During startup of the process recycle lines 14, 17, 21 and 22 are blocked off until sufficient product material is obtained for initiation of recycle. Similarly water diluent recycle through lines 13, 28 and 29 will be blocked off and instead an outside source of either water or steam will be injected into line 13 by means not shown just prior to the interconnection with line 31. At startup an alcoholic oxygenate feedstream flows via line 8 into zone 30 into contact with acidic porous alumina catalyst contained therein at etherification conditions effective to selectively convert at least about 10 to 70% or more of the alcoholic oxygenate to the corresponding ether. The etherification conditions in zone 30 include a temperature of 200 to 375° C, an inlet pressure of about 136 to 1136 kPa (5 to 150 psig) and a WHSV of about 0.5 to 5 hr-1. These conditions will convert from about 10 to 70% or more of methanol in the feedstream to dimethyl ether and water in an exothermic reaction that heats the ether-containing effluent stream 31 to a temperature from 375 to 525°C and shifts at least 10%, and preferably 20 to 30% or more, of the exothermic heat of reaction from the second reaction zone comprising moving bed reactors 1, 2 and 3 to this etherification step, thereby preheating the feedstream and avoiding detrimental temperature excursions in the second reaction zone. Line 31 withdraws the resulting heated ether-containing effluent stream from zone 30 via line 31 and suitable means (not shown) adjust the temperature of this stream to a range of 375 to 525° C to produce a heated mixture of Dimethyl ether, unreacted methanol and steam diluent. The Dimethyl ether synthesis reaction form steam diluent as by-product water in an amount of at least 0.5 mole per mole of methanol converted therein. The resulting heated mixture flows via line 31 to the intersection with line 13 that admixes an appropriate amount of diluent in the form of water or steam to provide an oxygenate to diluent ratio of about 0.5 :1 to about 5 :1 with a value of about 0.5:1 to 2:1 being especially preferred for startup. The resulting mixture of oxygenate feed and diluent passes through appropriate feed and effluent heat exchange and heating means (not shown) to fully vaporize the resulting stream and provide a charge stream for reactor 1 that enters the reactor at a temperature of about 375[deg.] to 525° C and a total pressure of about 136 to 1136 kPa (5 to 35 psig). Reactors 1, 2, and 3 contain sufficient dual- function catalyst to provide a Weight Hourly Space Velocity (WHSV) of about 0.5 to 5 hr-1 during startup with the effective WHSV rising to a level of about 1 to 10 hr-1 once recycle of olefinic hydrocarbon material commences. Line 9 withdraws effluent material from reactor 1, reduces its temperature to a value close to the temperature of the charge to reactor 1 via one or more cooling means not shown, and charges the resulting cooled effluent stream to reactor 2 wherein it contacts a dual-function catalyst to convert additional amounts of oxygenate material to propylene with production of an effluent stream withdrawn via line 10. Appropriate means (not shown) cool the effluent in line 10 to a temperature close to the inlet temperature to reactor 1 which passes into reactor 3 and contacts an additional amount of dual-function catalyst under conditions which result in further conversion of unreacted oxygenate to propylene and various other by-products. Maintaining similar temperature differentials across the reactor of the second reaction zone minimizes coking of the catalyst in reactors 2 and 3 to the same degree as in reactor 1. Line 11 withdraws the effluent stream from reactor 3, an appropriate cooling step liquefies a substantial amount of the water contained therein via one or more cooling means such as a feed/effluent exchanger (not shown), and passes the effluent to a three-phase separator, zone 5, in which forms a hydrocarbon vapor phase, a liquid hydrocarbon phase and a water phase containing significant amounts of any unreacted oxygenate that escapes from reactor 3. By maintaining the activity of the dual-function catalyst at essentially start-of-cycle conditions this invention expects to pass a minimal amount of unreacted oxygenate into zone 5 and to achieve an overall conversion of oxygenate feed achieving through 1, 2 and 3 of 97% or greater during the entire cycle. Line 13 withdraws the aqueous phase from separator 5. A drag stream taken by line 23 dispose of surplus water. Line 13 recycles the resulting net water stream into admixture with the oxygenate feed via line 31. Lines 28 and 29 may add additional quantities of water for cooling purposes to reactors 2 and 3. The initiation of water recycle terminates the startup provisions for injection of diluent water. A line 12 withdraws a vapor phase from separator 5 and charges it to fractionation column 6 which acts as a deethanizer and operates to produce an ethylene-rich overhead fraction 14 that also contains minor amounts of ethane and some methane and a bottom fraction 15 which essentially comprises the C3+ portion of the material charged to column 6. A line 18 withdraws a drag stream from the overhead stream 14 to control the buildup of C1 and C2 paraffins in the ethylene recycle loop. The ethylene recycle loop may require additional treatment of the ethylene rich overhead stream to remove methane from this stream and prevent the buildup of a significant methane concentration. Any suitable means may be used and include a demethanizer column, a methane adsorption zone, and a methane-selective membrane zone. Line 18 may take 1 to 15 vol-% of the overhead stream flowing through line 14 and more typically will comprise about 1 to 10 vol-% of this stream. Line 14 charges the remainder of the overhead stream to reactor 1 as an ethylene-rich recycle stream via line 31. The present invention may apportion the ethylene-rich recycle stream flowing in line 14 between the three reactor zones but the drawing depicts the preferred flow arrangement that charges the entire amount of this ethylene to reactor 1 and results in superior ethylene conversion when by its exposure to the maximum amount of available catalyst as it flows through the three reactor zones. Line 15 charges the C3- material of the bottom stream to depropanizer column 7 at fractionation conditions that produce a propylene-rich overhead stream 16 as the principal product stream of the present invention and which contains minor amounts of by-product propane materials. Line 17 withdraws the bottom stream of C4+ olefin-rich material which primarily comprises C4, C5 and C6 olefinic material along with very minor amounts of butane and pentane from column 7. A line 19 takes a drag stream in an amount of about 1 to 15 vol-% and preferably about 1 to 3 vol-% to control buildup of these paraffinic materials in this C4+ olefin recycle loop. Line 17 together with lines 22 and 23 pass the remainder of the bottom material to reactors 1, 2 and 3 to provide C4+ olefin reactants unblocking of lines21 and 22 occurs once recycle operations commence. Preferably the major portion of the bottom material from column 7 passes to reactor 1 and minor portions pass to reactors 2 and 3 with preferably 60 to 80 vol-% of the total amount of this C44 olefin rich recycle stream flowing to reactor 1 via lines 17 and equal portions of about 10 to 20 vol-% being simultaneously entering reactors 2 and 3 via lines 17, 21 and 9 for reactor 2 and via lines 17, 21, 22 and 10 for reactor 3. Once the three recycle streams flowing through lines 13, 14 and 17 are up and running then the startup mode of the instant invention is terminated and the full recycle mode of operation commences with the dual-function catalyst located in reactors 1, 2 and 3 functioning not only as oxygenate conversion catalyst but also as olefin interconversion catalyst. In the full recycle mode an alcoholic oxygenate rich feedstream enters the process via line 8 and is converted at least in part to the corresponding ether in zone 30. The resulting ether-containing mixture flowing via line 31 admixes with a first ethlylene rich recycle stream carried in line 14 and then with a water diluent in line 13. The resulting admixture passes to the intersection of line 31 where mixture with an additional quantity of C4+ olefinic material from line 17 forms the charge to reactor 1. This charge mixture flows into reactor 1 in a manner previously described after suitable heating to the appropriate inlet temperature. After passage through the bed of catalyst maintained in reactor 1 line withdraws the resulting effluent. Lines 28 adds an additional quantity of relatively cool water diluent along and line 21 adds an additional quantity of the relatively cool C4+olefin-rich recycle stream. After suitable cooling to achieve the prescribed inlet temperature for reactor 2 set forth the resulting mixture enters reactor 2 wherein it makes a passage through the bed of catalyst to produce an effluent stream. Line 10 withdraws the effluent and flows into admixture with an additional quantum of relatively cool water diluent which from line 29 and an additional quantum of relatively cool C4+ olefin-rich material from line 22. After cooling of the resulting mixture to the prescribed inlet temperature line 10 delivers the mixture to reactor 3 for contact with the catalyst therein and production of an effluent stream which after appropriate quench and cooling flows through line 11 to three-phase separating zone 5. The amount of dual-function catalyst utilized in reactors 1, 2 and 3 can vary. Reactors 2 and 3 may contain larger amounts of catalyst to make up for the slight amount of deactivation that occurs in reactor 1 when the catalyst flows to reactor 2 via line 25 and in reactor 2 when the catalyst flows from reactor 2 to reactor 3 via line 26. Preferably reactors 1, 2, and 3 run with essentially equal amounts of catalyst in the three zones or with a division which of approximately 25 to 30% of the total catalyst in each of reactors 1 and 2 and the remaining 40 to 50 vol-% in reactor 3. The OTP process starts catalyst circulation after the reactors are lined out at operating conditions. Catalyst circulates between the reactors via lines 25 and 26 and to regenerator zone 4 via line 27. Regeneration zone 4 removes coke deposits from the coke- containing catalyst charged thereto via line 27 using a low severity oxidative procedure as previously explained. The regeneration zone 4 also receives by means not shown an oxygen- containing gas stream containing about 0.5 to 2 vol-% oxygen which and supplied in an amount sufficient to support combustion of a major portion of the coke charged to this zone. Line 24 completes a catalyst circulation circuit defined by lines 25, 26, 27 and 24 by recirculating catalyst to reactor 1 via. A flue gas stream is withdrawn from zone 4 via a line not shown. The flow of catalyst around this catalyst circulation circuit is selected to provide a catalyst on-stream cycle time of 300 hours or less to maintain catalyst activity, oxygenate conversion and propylene selectivity at or near start of cycle conditions. Accordingly the catalyst particles flow around this circuit such that a catalyst particle's residence time in reactors 1, 2 and 3 is not more than 300 hours before it returns to zone 4 for regeneration. Line 32 withdraws the hydrocarbon phase from three-phase separator zone 5. This material generally boils in the gasoline range and can comprise a gasoline product stream of the present invention which may require further treatment due to the high content of olefinic material that is present therein. Preferably the liquid hydrocarbon phase of line 32 passes to an additional fractionation step (not shown) to recover a C4 to Cg olefin-rich overhead stream which can return to reactors 1, 2 and 3 to provide additional conversion of heavy olefins to propylene. Recycling of any such C4 to C6 olefin-rich stream may occur by directly charging it to column 7 or admixing it with the stream flowing in line 17. We claim: 1. A continuous process for selective conversion of an alcoholic oxygenate feed to propylene comprising the steps of: a) contacting the feed with an acidic etherification catalyst in a first reaction zone at etherification conditions effective to form an ether-containing effluent stream and to shift at least 10% of the exothermic heat of reaction liberated when the feed is converted to propylene from the subsequent propylene synthesis step to this etherification step, thereby heating the ether-containing effluent stream to a temperature of 250 to 450°C and producing by-product water in an amount of at least 0.5 mole per mole of alcoholic oxygenate converted; b) adjusting the temperature of the resulting ether-containing effluent stream to a range of 375 to 525°C and adding diluent thereto to produce a heated mixture of ether, unreacted alcoholic oxygenate and diluent; c) reacting the resulting heated mixture with dual-function catalyst particles containing a molecular sieve, having the ability to convert the oxygenates contained therein to C3 olefin and to interconvert C2 and C4+ olefins to C3 olefin, in a second reaction zone containing at least one moving bed reactor wherein the reaction zone is operated at oxygenate conversion conditions effective to convert the oxygenates contained in the mixture to propylene and at a catalyst circulation rate through the second reaction zone selected to result in a catalyst on-stream cycle time of 300 hours or less to produce a propylene-containing effluent stream containing major amounts of a C3 olefin product and water, lesser amounts of a C2 olefin, C4+ olefins and C1 to C4+ saturated hydrocarbons and minor amounts of unreacted oxygenate, by-product oxygenates and aromatic hydrocarbons; d) passing the propylene-containing effluent stream to a separation zone and therein cooling and separating this effluent stream into a vaporous fraction rich in C3 olefin, a water fraction containing unreacted oxygenates and by-product oxygenates and a liquid hydrocarbon fraction containing heavier olefins, heavier saturated hydrocarbons and minor amounts of aromatic hydrocarbons; e) recycling at least a portion of the water fraction recovered in step d) to step b) to provide a portion of the diluent added therein; f) separating the vaporous fraction into a C2 olefin-rich fraction, a C3 olefin-rich product fraction and C4+ olefin-rich fraction; g) recycling a portion of the C2 olefin-rich fraction or of the C4+ olefin-rich fraction or of a mixture of these fractions to step c); and h) withdrawing coke-containing dual-function catalyst particles from the second reaction zone, oxidatively regenerating the withdrawn catalyst particles in a regeneration zone and returning a stream of regenerated catalyst particles to the second reaction zone. 2. The process as claimed in claim 1 wherein the alcoholic oxygenate is an aliphatic alcohol containing 1 to 6 carbon atoms. 3. The process as claimed in claim 2 wherein the aliphatic alcohol is methanol the acidic etherification catalyst is an alumina catalyst and first reaction zone produces a -dimethyl ether (DME) containing effluent. 4. The process as claimed in claims 1,2 or 3 wherein the etherification conditions include a temperature of 200 to 375°C, an inlet pressure of about 136 to 1136 kPa (5 to 150 psig) and a weight hour space velocity (WHSV) of 0.1 to 10 hr'l. 5. The process as claimed in claims 1,2, or 3 wherein the dual function catalyst contains a zeolitic molecular sieve and or an ELAPO molecular sieve. 6. The process as claimed in claim 5 wherein the zeolitic molecular sieve has a structure corresponding to ZSM-5 and the ELAPO molecular sieve is a SAPO material having a structure corresponding to SAPO-34. 7. The process as claimed hi claims 1,_2, or 3 wherein the second reaction zone contains at least three moving bed reactors, the moving bed reactors are connected in a serial flow configuration with respect to the stream of catalyst particles that passes there through, and the moving bed reactors are connected in a serial flow configuration with respect to the flow of the heated mixture fed thereto. 8. The process as claimed in claim 1,_2, or 3 wherein the liquid hydrocarbon fraction separated in step d) is further separated into a C4 to C6 olefin-rich fraction and a naphtha product fraction and least a portion of the C4 to C6 olefin-rich fraction is recycled to step c). 9. The process as claimed in claim 8 wherein a portion of the C2 olefin-rich fraction, the C4+ olefin-rich fraction and the C4 to C6 olefin-rich fraction are recycled to step c). 10. The process as claimed in claim 1,_2, or 3 wherein the dual-function catalyst particles are regenerated in the regeneration zone using an oxygen-containing stream under conditions selected to produce a regenerated catalyst containing less than 0.5 wt-% carbonaceous material. 11. A continuous process for selective conversion of an alcoholic oxygenate feed to propylene, substantially as hereinbefore described with reference to the accompanying drawings.

Full Text

BACKGROUND OF THE INVENTION
The present invention relates generally to the use of an etherification reaction in a pretreatment step performed on the alcoholic oxygenate feed and of moving bed technology in the hydrocarbon synthesis reaction zone of an oxygenate to propylene (OTP) process. These features are coupled with a relatively short OTP catalyst on-stream cycle time to result in a level of OTP catalyst coking that does not significantly degrade the activity and propylene selectivity of the dual-function catalyst used in this OTP process. These three provisions sharply improve the average propylene yield achievable by this process over its on-stream catalyst cycle time relative to the average cycle propylene yield that is achievable by the prior art process that uses fixed bed technology for the oxygenate to hydrocarbon conversion reaction and operates with much longer catalyst cycle times.
A major portion of the worldwide petrochemical industry is concerned with the production of light olefin materials and their subsequent use in the production of numerous important chemical products via polymerization, oligomerization, alkylation and the like well-known chemical reactions. Light olefins include ethylene, propylene and mixtures thereof. The art has long sought a source other than petroleum for the massive quantities of raw materials that are needed to supply the demand for these light olefin materials. Oxygenates are particularly attractive alternate source because they can be produced from such widely available materials as coal, natural gas, recycled plastics, various carbon waste streams from industry and various products and by-products from the agricultural industry. The art of making methanol and other oxygenates from these types of raw materials is well established. The prior art has revealed essentially two major techniques that are discussed for conversion of methanol to light olefins (MTO). The first of these MTO processes is represented in US-A-4,387,263. US-A-4,587,373 discloses the need to operate at a substantial pressure to make the commercial equipment of reasonable size and the diversion of a portion of the methanol feed to the dimethyl ether absorption zone to downsize the scrubbing zone.
To control the amounts of undesired C4+ hydrocarbon products produced by ZSM-5 type of catalyst systems later prior art uses a non-zeolitic molecular sieve catalytic material. US-A-5,095,163; US-A-5,126,308 and US-A-5,191,141 disclose a metal aluminophosphate (ELAPO) and more specifically a silicoaluminophosphate molecular sieve (SAPO), with a strong preference for SAPO-34. The classical oxygenate to olefin (OTO) technology produces a mixture of light olefins primarily
ethylene and propylene along with various higher boiling olefins. Although the classical OTO process technology possesses the capability of shifting the major olefin product recovered therefrom from ethylene to propylene by various adjustments of conditions maintained in the reaction zone, the art has long sought an OTP technology that would provide better yields of propylene relative to the classical OTO technology.
Publication No. US2003/0139635A1 which was published on July 24, 2003 and describes a process for selectively converting methanol to propylene (MTP) and/or converting dimethyl ether to propylene and utilizes three reactors containing fixed beds of oxygenate conversion catalysts in a parallel flow arrangement with respect to the oxygenate feed and a serial flow arrangement with respect to the effluents of the first reactor and the second reactor. The publication teaches a dual function OTP catalyst of a pentasil-type (i.e. ZSM-5 type) having an alkali content less than 380 ppm and a zinc oxide content of less than 0.1 wt- % coupled with a restriction on cadmium oxide content of the same amount. This MTP process is further described in Rothaemel et al. "Demonstrating the New Methanol to Propylene (MTP) Process" (presented to the ERTC Petrochemical Conference in March of 2003 at Paris, France) as having an expected on-stream portion of the process cycle of 500 to 700 hours before in situ regeneration becomes necessary and shows a significant drop in conversion activity over the first five cycles. The conventional procedure for compensating for activity decay in a catalytic operation involves increasing the average reactor temperature to attempt to hold conversion in the targeted range of greater than 94% of the oxygenate charge.
The problem addressed by the present invention is to enhance the average propylene selectivity of an OTP process over its on-stream cycle time and thereby diminish the requirement for recycle of olefin products other than propylene in order to compensate for lower propylene selectivity.
SUMMARY OF THE INVENTION
We have now discerned that propylene selectivity is a function not only of reaction conditions but also of average coke level deposited on the OTP conversion catalyst during the on-stream portion of the process cycle. Coke deposition is accelerated when the AT (i.e. T at the outlet of the reactor minus T at the inlet) across the individual OTP reaction is not carefully controlled. This is caused by the strongly exothermic nature of the desired OTP reactions which at 400 to 500° C can reach levels of 7 to 11 kcal/mole of oxygenate converted and which can cause the temperature at the tail-end of the reactor to
reach levels that accelerate coke deposition and thus significantly degrade propylene selectivity. We have now found that average propylene selectivity in an OTP process can be significantly enhanced if the amount of detrimental carbonaceous deposits laid down on the catalyst during the on-stream portion of the process cycle is held to a level such that the OTP catalyst activity is not significantly diminished and oxygenate conversion is maintained over the process cycle near or at essentially start of cycle level. By the use of the term "not significantly diminished" we intend to mean that the conversion level at constant conditions is not allowed to fall more than 2 to 3% during the course of the on-stream catalyst cycle. Quite surprisingly we have also discerned that if this restriction is put on such an OTP process then the average propylene selectivity over the cycle will be improved by 1.5% to 5.5% or more. In other words if the once-through propylene selectivity of the prior art OTP process is in the range of 30% to 40% of the carbon-containing products from the OTP reactions we would anticipate the use of our invention to enable an improvement in this range to a level of at least 31.5% to 41.5% or more, thereby significantly improving the economics of the process. In sharp contrast to the OTP process of the prior art, our invention envisions shifting at least 10%, and preferably 20 to 30% or more, of the exothermic heat of reaction liberated when an alcoholic oxygenate such as methanol is converted to propylene from the OTO conversion step to a feed pretreatment step involving an etherification reaction which is designed to convert alcoholic oxygenates such as methanol to their corresponding ether (i.e. Dimethyl ether in the case of methanol) with resulting release of energy that is used to preheat and/or vaporize the feed rather than to accelerate coke deposition in the individual OTP reactors. This strategy allows tighter control of the AT across each of the OTP reactors and thus results in less detrimental coke deposition on the OTP catalyst and in lower permanent deactivation due to steam dealumination which accelerates at higher reactor temperatures. When this [Delta]T control strategy is coupled with our concept of replacing the fixed bed technology in the olefin synthesis reactors of the prior art with moving bed technology, it provides an OTP process that meets the target of not significantly disturbing initial OTP catalyst activity and propylene selectivity while operating with a catalyst circulation rate through the reactors such that the catalyst on-stream cycle time is 300 hours or less. Under these conditions the initial on- stream alcoholic oxygenate conversion of the OTP process does not drop by more than 2 to 3% thereby enabling maintenance of propylene selectivity at essentially start of run condition.
The primary objective of the present invention is to provide a realistic and technically feasible solution to the problem of propylene selectivity loss during the on-stream cycle of the prior art fixed bed OTP
process. A secondary objective is to improve the economics of the prior art OTP process in order to diminish the amount of olefins other than propylene that must be recycled in order to maintain overall propylene selectivity at the target level. A third objective is to control OTP reactor temperature differential by shifting at least 10% of the exothermic heat of reaction liberated when an alcoholic oxygenate such as methanol is converted to propylene from the OTP reaction step to a feed pretreatment step that catalytically etherifies feed alcoholic oxygenates to selectively produce dimethyl ether and the like ethers. Another object of the present invention is to avoid severe deactivation of the dual-function OTP catalyst utilized in the OTP process of the present invention in order to minimize the severity of the regeneration step that is needed in order to restore fresh catalyst activity, thereby minimizing hydrothermal damage and prolonging catalyst life.
In one embodiment the instant invention is a continuous process for the selective conversion of an alcoholic oxygenate feed to propylene. In the first step of the process an oxygenate feed comprising alcohol contacts an acidic etherification catalyst in a first reaction zone at etherification conditions effective to form an ether-containing effluent stream and to shift at least 10% of the exothermic heat of reaction liberated when the feed is converted to propylene from the subsequent propylene synthesis step to this etherification step, thereby heating the ether-containing effluent stream to a temperature from 250 to 45O0C and producing by-product water in an amount of at least 0.5 moles per mole of alcoholic oxygenate converted. The second step adjusts the temperature of the resulting ether-containing effluent stream to a temperature from 375 to 525° C and adds sufficient diluent thereto to produce a heated mixture of ether, unreacted alcoholic oxygenate and diluent. The resulting heated mixture contacts a dual-function catalyst containing a molecular sieve known to have the ability to convert at least a portion of the oxygenates contained therein to propylene and to interconvert C2 and C4+ olefins to C3 olefins. This OTP reaction step is performed in a second reaction zone containing at least one moving bed reactor which operates at oxygenate conversion conditions effective to convert oxygenates contained in the heated mixture to propylene and at a catalyst circulation rate through the second reaction zone selected to result in an OTP catalyst on-stream cycle time of 300 hours or less. An effluent stream is then withdrawn from the second reaction zone that contains major amounts of a C3 olefin product and a water by-product and lesser amounts of a C2 olefin, C4+ olefins, C1 to C4+ saturated hydrocarbons and minor amounts of unreacted oxygenate, by-product oxygenates and aromatic hydrocarbons. In the next step this effluent stream passes to a separation zone, is cooled therein and separated into a vaporous fraction rich in C3 olefins, a water fraction containing unreacted
oxygenate and by-product oxygenates and a liquid hydrocarbon fraction containing heavier olefins, heavier saturated hydrocarbons and minor amounts of aromatic hydrocarbons. At least a portion of the water fraction recovered in this separation step is then recycled to the second step to provide at least a portion of the diluent used therein. The vaporous fraction recovered in this separation step is further separated in a second separating zone into a C2 olefin-rich fraction, a C3 olefin-rich product fraction and a C4+ olefin-rich fraction. The C3 olefin-rich product fraction is then recovered as a principal product stream from the present process and at least a portion of the C2 olefin-rich fraction or of the C4+olefin-rich fraction or of a mixture of these fractions is recycled to the propylene synthesis step in order to interconvert these materials into additional quantities of the desired propylene product. In another step of the process a stream of coke-containing dual-function catalyst particles is withdrawn from the second reaction zone, oxidatively regenerated with an oxygen-containing stream in a regeneration zone and a stream of regenerated catalyst particles is returned to the second reaction zone in order to provide regenerated catalyst for circulation therethrough.
A second embodiment involves a continuous process for the selective conversion of an alcoholic oxygenate feed to propylene as described in the first embodiment wherein the dual-function catalyst contains a zeolitic molecular sieve having a structure corresponding to ZSM-5, or an ELAPO molecular sieve having a structure corresponding to SAPO-34 or a mixture of these materials.
Another embodiment comprises a continuous process for selective conversion of an alcoholic oxygenate feed to propylene as described above in the first embodiment wherein the second reaction zone contains at least 3 moving bed reactors which are connected in a serial flow or parallel flow configuration with respect to heated mixture fed thereto and in a serial flow configuration with respect to the stream of catalyst particles that passes therethrough.
A highly preferred embodiment of the present invention comprises a continuous process for the selective conversion of an alcoholic oxygenate feed to propylene as described above in the first embodiment wherein the alcoholic oxygenate feed essentially comprises methanol.
A high propylene yield embodiment of the instant process involves the continuous process for selective conversion of an alcoholic oxygenate feed to propylene as described in any of the previous embodiments wherein the liquid hydrocarbon fraction recovered in the first separation step is further separated into a C4 to C6 olefin-rich fraction and a naphtha product fraction and at least a portion of
the C4 to C6 olefin-rich fraction is recycled to the OTP conversion step in order to interconvert these heavier olefins into propylene.
BRIEF DESCRIPTION OF THE DRAWING
The FIGURE is a process flow diagram of a preferred embodiment of the present invention.
TERMS AND CONDITIONS DEFINITIONS
The following terms and conditions are used in the present specification with the following meanings: (1) A "portion" of a stream means either an aliquot part that has the same composition as the whole stream or a part that is obtained by eliminating one or more readily separable component therefrom. (2) An "overhead" stream means the net overhead recovered from the specified zone after recycle of any portion to the zone for reflux or any other reason. (3) A "bottom" stream means the net bottom stream from the specified zone obtained after recycle of any portion for purposes of reheating and/or reboiling and/or after any phase separation. (4) The term "light olefins" means ethylene, propylene and mixtures thereof. (5) The expression "OTP" process means a process for converting an alcoholic oxygenate to propylene and in a preferred embodiment when the alcoholic oxygenate is methanol the OTP process is referred to as an MTP process herein. (6) The term "catalyst on- stream cycle time" means the length of time the catalyst particle is exposed to feed at conversion conditions before withdrawal from the reaction zone for regeneration. (7) The term "propylene average cycle yield" means the total propylene yield during the catalyst on- stream cycle time divided by the total amount of oxygenate feed converted during the catalyst on-stream cycle time. (8) The term "dual-functional" means that the OTP catalyst catalyzes both the OTP reactions and the olefin interconversion reactions necessary to convert C2 and C4+ olefins to propylene.
STATEMENT OF THE INVENTION
The present invention relates to a continuous process for selective conversion of an alcoholic oxygenate feed to propylene comprising the steps of:
a) contacting the feed with an acidic etherification catalyst in a first reaction zone at
etherification conditions effective to form an ether-containing effluent stream and to shift at least 10% of the exothermic heat of reaction liberated when the feed is converted to propylene from the subsequent propylene synthesis step to this etherification step, thereby heating the ether-containing effluent stream to a temperature from 250 to 450° C and producing by-product water in an amount of at least 0.5 mole per mole of alcoholic oxygenate converted;
b) adjusting the temperature of the resulting ether-containing effluent stream to a range of 375
to 525° C and adding diluent thereto to produce a heated mixture of ether, unreacted alcoholic
oxygenate and diluent;
c) reacting the resulting heated mixture with dual-function catalyst particles containing a
molecular sieve, having the ability to convert the oxygenates contained therein to C3 olefin and
to inter convert C2 and C4+ olefins to C3 olefin, in a second reaction zone containing at least
one moving bed reactor wherein the reaction zone is operated at oxygenate conversion
conditions effective to convert the oxygenates contained in the mixture to propylene and at a
catalyst circulation rate through the second reaction zone selected to result in a catalyst on-
stream cycle time of 300 hours or less to produce a propylene-containing effluent stream
containing major amounts of a C3 olefin product and water, lesser amounts of a C2 olefin, C4+
olefins and C1 to C4+saturated hydrocarbons and minor amounts of unreacted oxygenate, by
product oxygenates and aromatic hydrocarbons;
d) passing the propylene-containing effluent stream to a separation zone and therein cooling and separating this effluent stream into a vaporous fraction rich in C3 olefin, a water fraction containing unreacted oxygenates and by-product oxygenates and a liquid hydrocarbon fraction containing heavier olefins, heavier saturated hydrocarbons and minor amounts of aromatic hydrocarbons;
e) recycling a portion of the water fraction recovered in step d) to step b) to provide a portion of the diluent added therein;
f) separating the vaporous fraction into a C2 olefin-rich fraction, a C3 olefin-rich product fraction and a C4+ olefin-rich fraction;
g) recycling a portion of the C2 olefin-rich fraction or of the C44 olefin-rich fraction or of a mixture of these fractions to step c); and
h) withdrawing coke-containing dual-function catalyst particles from the second reaction zone, oxidatively regenerating the withdrawn catalyst particles in a regeneration zone and returning a stream of regenerated catalyst particles to the second reaction zone.
DETAILED DESCRIPTION OF THE INVENTION
The feedstream charged to the instant OTP process comprises one or more alcoholic oxygenates in admixture with an aqueous diluent. The concentration of the one or more alcoholic oxygenates in the feedstream is preferably at least 70 wt-% while a concentration of 90%+ is preferred with best results obtained when the alcoholic oxygenate in the feedstream is at a level of 95 wt-% or more. The alcoholic oxygenate that is present in the feedstream preferably contains at least one oxygen atom and about 1 to 6 carbon atoms with best results usually obtained with an aliphatic alcoholic oxygenate containing 1 to 4 carbon atoms. Suitable alcoholic oxygenates for use in the instant process include saturated straight or branch chain aliphatic alcohols and their unsaturated counterparts. Representative of suitable aliphatic alcohols include methanol, ethanol, isopropyl alcohol, normal propyl alcohol, butyl alcohol, isobutyl alcohol, t-butyl alcohol, amyl alcohol and mixtures of these aliphatic alcohols. Methanol is the most preferred alcoholic oxygenate for use in the feed and particularly preferred results are obtained with a feedstream which contains a relatively high concentration of methanol. Commercial feedstreams of a so-called "crude" methanol stream usually contains 70 to 80 wt-% methanol in admixture with water and of a purified methanol stream typically contains methanol in a concentration of 95 wt-% or more in admixture with trace amounts of water.
In the first reaction step of the instant process the feedstream is contacted with an acidic etherification catalyst in a first reaction zone at etherification conditions effective to form an effluent stream containing at least one ether corresponding to the alcoholic oxygenate present in the feedstream. If the feedstream contains only methanol the etherification forms dimethyl ether and if the feedstream contains only ethanol diethylether (DEE) if formed with corresponding simple and mixed ethers formed in the case where the feedstream contains a mixture of alcoholic oxygenates. A significant feature of the invention is performing the etherification step at a conversion level sufficient to shift at least 10%, and preferably 20 to 30% or more, of the exothermic heat of reaction liberated when the feed is converted to propylene from the subsequent propylene synthesis step to this etherification step.
From a heat integration standpoint it is highly preferred to operate the etherification reaction to preheat and/or vaporize at least a portion of the feedstream and to cut the exothermicity associated with the downstream processing of the resulting ether-rich stream in the propylene synthesis (OTP) reaction zone. The reaction of an ether such as Dimethyl ether to produce propylene and other olefins is significantly less exothermic than the direct conversion of an alcohol to an olefin in the OTP reaction zone. Accordingly the feed pretreatment step of etherification captures a significant portion of the exothermic heat released in the overall conversion of an alcoholic oxygenate to the desired olefinic hydrocarbons, thereby providing a convenient source of heat for bringing the feedstream up to the desired reaction for the propylene synthesis step and/or vaporizing a significant portion of the feedstream to facilitate a vapor phase reaction in the olefin synthesis step. An unobvious benefit of this shifting of a significant portion of the heat released during the highly exothermic conversion of an alcohol to olefinic hydrocarbons is the enabling of pretreatment step to provide sharper control over the AT experienced in the olefin synthesis reaction step. Since the coke forming side reaction is accelerated at higher temperatures and the catalyst dealumination reaction is similarly accelerated the ability to control the temperature increase in the OTP reaction zone suppresses not only this coke formation side reaction that degrades high propylene selectivity but also this catalyst dealumination reaction. A significant benefit of the instant invention is thus that the selectivity of the OTP reaction to the desired propylene olefin is enhanced by the fact that tight control can be maintained over the average temperature increase experienced in the olefin synthesis reaction zones.
The feed pretreatment step or etherification step of the present invention is catalytic and requires the presence of mildly acidic heterogeneous catalyst in order to accelerate the reaction. Suitable catalysts for the etherification step of the present invention are mildly acidic porous solids having a surface area of about 10 to 500 m2/g and include the following materials: (1) silica or silica gel, clays and silicates including those synthetically prepared and naturally occurring, which may or may not be acid treated, for example, attapulgus clay, china clay, diatomaceous earth, fuller's earth, kaolin, bentonite, kieselgurh, etc. (2) refractory inorganic oxides such as alumina, titanium dioxide, zirconium dioxide, chromium oxide, beryllium oxide, vanadium oxide, cesium oxide, hafnium oxide, zinc oxide, magnesia, boria, thoria, silica-alumina, silica-magnesia, chromia-alumina, alumina-boria, silica-zirconia, etc. (3) crystalline zeolitic aluminosilicates such as naturally occurring or synthetically prepared mordenite and/or faujasite, ZSM-5, ZSM-11, or ZSM-12, either in the hydrogen form or in a
form which has been treated with multivalent cations; and (4) combinations of materials from one or more of these groups. The preferred porous catalysts for use in the present invention are refractory inorganic oxides, with best results obtained with an alumina material.
Suitable alumina materials are the crystalline aluminas known as gamma-, eta- and theta-alumina, with gamma- or eta-alumina giving best results. In addition, in some embodiments, the alumina carrier material may contain minor proportions of other well known refractory inorganic oxides such as silica, zirconia, magnesia, etc.; however, the preferred catalyst is substantially pure gamma- or eta-alumina. Preferred catalysts have an apparent bulk density of about 0.3 to about 0.9 g/cc and surface area characteristics such that the average pore diameter is about 20 to 300 Angstroms, the pore volume is about 0.1 to about 1 cc/g and the surface area is about 100 to about 500 m2/g. In general, best results are typically obtained with a gamma-alumina carrier material which is used in the form of spherical particles having: a relatively small diameter (i.e. typically about 1/16 inch), an apparent bulk density of about 0.3 to about 0.8 g/cc, a pore volume of about 0.4 ml/g, and a surface area of about 150 to about 250 m2/g.
The etherification conversion conditions utilized in the present invention include a temperature from 200 to 375°C with a preferred range of 275 to 35O0C, an inlet pressure from 136 to 1136 kPa (5 to 150 psig) with best results achieved at a pressure from 136 to 791 kPa (5 to 100 psig). For the lower boiling alcoholic oxygenates such as methanol and ethanol the temperature and pressure used in the etherification step are usually chosen to maintain substantial vapor phase conditions or a mixed liquid/vapor condition. The weight hourly space velocity (WHSV) which is utilized in the etherification step is calculated on the basis of the mass hourly flow rate of the sum of the mass of the alcoholic oxygenate material contained in the feedstream plus any alcoholic oxygenate material contained in recycle streams divided by the mass of the etherification catalyst present in the etherification conversion zone. Ordinarily good results are obtained with WHSV of about 0.1 to 10 hrs-1 with best results ordinarily achieved in the range of about 0.5 to 5 hrs-1. In accordance with the present invention these conversion conditions are adjusted to achieve a release of at least 10%, and preferably about 20 to 30% or more, of the exothermic heat of reaction liberated when the alcoholic oxygenate contained in the feed is converted to propylene and other olefins as previously explained.
As a result of the exothermic heat liberated in the etherification step the effluent stream withdrawn therefrom is recovered at a temperature from 250 to 45O0C and contains by-product water in amounts
of at least 0.5 moles per mole of alcoholic oxygenate converted in the etherification step. This effluent stream from the etherification step also contains substantial amounts of one or more ethers corresponding to the one or more alcoholic oxygenates present in the feedstream. For example, with preferred alcoholic oxygenate of methanol the effluent stream from this step will contain substantial quantities of dimethyl ether.
In the next step of the present invention the temperature of the resulting ether- containing effluent stream is adjusted to a range from 375 to 525° C either by adding or subtracting the necessary thermal energy from this effluent stream utilizing suitable heating and/or cooling means well known to those skilled in the art. It is also a preferred practice during this effluent temperature adjustment step to add sufficient diluent to this effluent stream to result in a mole ratio of diluent to oxygenate equivalents (i.e. the unreacted alcoholic oxygenates plus the ether content of the effluent stream) of about 0.5:1 to 5:1 with best results usually achieved at a diluent ratio of about 1:5 to 2:1. It is of course within the scope of the present invention to perform this diluent adjustment step at least in part after the ether-containing effluent stream has been passed through one or more of the downstream propylene synthesis reaction zones in the case where a multibed propylene synthesis scheme like that shown in the attached drawing is employed. After completing adjustments to the temperature and diluent content of the effluent stream withdrawn from the etherification step the resulting heated mixture of ether, unreacted alcoholic oxygenate and diluent passes to the propylene synthesis reaction step of the present invention (the OTP conversion step) to selectively prepare the desired propylene product.
In the OTP conversion step of the present invention, the resulting heated mixture is catalytically converted to hydrocarbons containing aliphatic moieties such as - but not limited to - methane, ethane, ethylene, propane, propylene, butylene, and limited amounts of other higher aliphatics by contacting the heated mixture with a dual-function OTP catalyst. Presence of a diluent is useful to maintain the selectivity of the OTP catalyst to produce light olefins, particularly propylene. The use of a diluent such as steam can provide certain equipment cost and thermal efficiency advantages as well as lowering the partial pressure of the oxygenate reactants, thereby increasing selectivity to olefins. The phase change between steam and liquid water can also advantageously transfer heat between the feed to the OTP reaction zone and the OTP reactor effluent, and the separation of the steam diluent from the product requires only a simple condensation step to separate water from the hydrocarbon products. [0024] As previously stated, the amount of diluent used in the OTP reaction step will preferably the range from
about 0.5:1 to 5:1 moles of diluent per mole of oxygenate equivalent and preferably 0.5:1 to 2:1 in order to lower the partial pressure of the oxygenates to a level which favors production of propylene. Some of the embodiments of the present invention envision recycling of certain olefinic streams that contain significant amounts of olefins other than propylene and saturated hydrocarbons. These recycle streams will furnish diluent to the OTP reaction zone and may reduce the amount of diluent that must be added to the OTP reaction zone to achieve the target diluent to oxygenate mole ratio once the OTP reaction zone is started. In the most preferred case water is used as the primary diluent and the amount of water charged to the OTP reaction zone during startup will diminish in proportion to the amount of other diluents that are recycled to this reaction zone.
The conversion conditions used in the OTP reaction zone are carefully chosen to favor the production of propylene from the oxygenates. An oxygenate conversion temperature range of 350° to 600° C is effective for the conversion of oxygenate over the known oxygenate conversion catalysts. The lower temperature in this range favor the production of propylene and upper temperature of the range favor the production of ethylene. Preferred inlet temperatures to the OTP reaction zone range from 375° to 525° C and more preferably from 400° C to 500° C. The temperature increase across each of the OTP reactors is preferably held to a range of 10° to 60° C to minimize hydrothermal deactivation and to avoid the acceleration of coke deposition on the catalyst that occurs when the tail end temperature individual reactors builds to levels beyond those contemplated by the present invention. Most methods for controlling the temperature increase in a reaction zone involve utilization of multiple beds of catalyst in separate reactors with inner-bed or inter-bed cooling utilizing appropriate heat exchange and/or addition of relatively cool recycle streams, or portions of the oxygenate feed and/or the diluents that is utilized in the zone. The present invention contemplates the use of lighter and heavier olefin interconversion reactions which are mildly endothermic to help control reactor temperature increases to the specified range. The preferred mode of operation of the instant invention uses at least 3 moving bed reactors with interbed quench achieved at least in part by the use of relatively cool recycle streams to provide additional quantities of reactants and diluent.
The oxygenate to propylene conversion step is effectively carried out over a wide range of pressures including inlet total pressures of about 0.1 to 100 atm (10.1 kPa to 10.1 mPa) and more typically between about 136 to 1136 kPa (5 to 150 psig). Since the formation of lighter olefins like propylene are
favored at low pressure conditions the preferred inlet pressure ranges from 136 to 343 kPa (5 to 35 psig).
The contact time of the reactants with the dual-function catalyst is ordinarily measured in relative terms of a Weight Hourly Space Velocity (WHSV) calculated on the basis of a mass hourly flow rate of the sum of the mass of oxygenate reactants passed to the OTP conversion zone plus the mass of any reactive hydrocarbon material present in the f eedstream or any of the recycle streams divided by the mass of the dual-function catalyst present in the OTP conversion zone. Weight Hourly Space Velocity (WHSV) in the OTP conversion zone ranges from about 0.1 to 100 hr-1, with a preferred range of 0.5 to 20 hr-1, and with best results ordinarily attained in the range of 0.5 to 10 hr-1.
In the oxygenate to propylene conversion step it is preferred to use a dual-function catalyst having the capability of converting oxygenates to propylene as well as the capability of interconverting olefins other than propylene to propylene. Any known catalytic materials known having the capability to catalyze these two reactions are suitable catalysts for the present invention. The preferred dual-function catalyst contains a molecular sieve as the active ingredient and more specifically the molecular sieve has relatively small pores. The preferred small pore catalysts are defined as having pores at least a portion, preferably a major portion, of which have an average effective diameter characterized such that the adsorption capacity (as measured by the standard McBain-Bakr gravimetric adsorption method using given adsorbate molecules) shows good adsorption of oxygen (average kinetic diameter of about 0.346 run) and negligible adsorption of isobutane (average kinetic diameter of about 0.5 nm). More preferably the average effective diameter is characterized by good adsorption of xenon (average kinetic diameter of about 0.4 nm) and negligible adsorption of isobutane, and most preferably, by good adsorption of n-hexane (average kinetic diameter of about 0.43 nm) and negligible adsorption of isobutane. Negligible adsorption of a given adsorbate is adsorption of less than three percent by weight of the catalyst and good adsorption of the adsorbate is over three percent by weight of the adsorbate based on the weight of the catalyst. Certain of the catalysts useful in the present invention have pores with an average effective diameter of less than 5 A. The average effective diameter of the pores of preferred catalysts is determined by measurements described in D. W. Breck, ZEOLITE MOLECULAR SIEVES by John Wiley & Sons, New York (1974). The term "effective diameter" denotes that occasionally the pores are irregularly shaped, e.g., elliptical, and thus the pore dimensions
are characterized by the molecules that can be adsorbed rather than the actual dimensions. Preferably, the small pore catalysts have a substantially uniform pore structure. Suitable dual-function catalysts may be chosen from among zeolitic molecular sieves and non-zeolitic molecular sieves.
Zeolitic molecular sieves in the calcined form may be represented by the general formula:
(Formula Removed)
where Me is a cation, x has a value from about 2 to infinity, n is the cation valence and y has a value of about 2 to 100 or more and more typically about 2 to 25.
Zeolites which may be used include chabazite - also referred to as Zeolite D, clinoptilolite, erionite, ferrierite, mordenite, Zeolite A, Zeolite P, ZSM-5, ZSM-11, and MCM- 22. Zeolites having a high silica content, (i.e., Those having silica to alumina ratios greater than 10 and typically greater than 100 with best results achieved at a silica to alumina mole ratio of about 250:1 to 1000:1 are especially preferred.) One such high silica zeolite having the structure of ZSM-5 is silicalite, as the term used herein includes both the silicapolymorph disclosed in US-A-4,061,724 and also the F-silicate disclosed in US-A-4,073,865.
Non-zeolitic molecular sieves include molecular sieves which have the proper effective pore size and are embraced by an empirical chemical composition, on an anhydrous basis, expressed by the empirical formula:
(Formula Removed)
where EL is an element selected from the group consisting of silicon, magnesium, zinc, iron, cobalt, nickel, manganese, chromium and mixtures thereof, x is the mole fraction of EL and is at least 0.005, y is the mole fraction of aluminum and is at least 0.01, z is the mole fraction of phosphorous and is at least 0.01 and x + y + z = 1. When EL is a mixture of metals, x represents the total amount of the element mixture present. Preferred elements (EL) are silicon, magnesium and cobalt with silicon being especially preferred.
The preparation of various ELAPOs may be found in US-A-5,191,141 (ELAPO); US-A-4,554,143 (FeAPO); US-A-4,440,871 (SAPO); US-A-4,853,197 (MAP05 MnAPO, ZnAPO, CoAPO); US-A-4,793,984 (CAPO); US-A-4,752,651 and US-A-4,310,440.
A preferred embodiment of the invention is one in which the element (EL) content varies from about 0.005 to about 0.05 mole fraction. If EL is more than one element, then the total concentration of all the elements is between about 0.005 and 0.05 mole fraction. An especially preferred embodiment is one in which EL is silicon (usually referred to as SAPO). The SAPOs which can be used in the instant invention are any of those described in US-A- 4,440,871; US-A-5,126,308, and US-A-5,191,141. SAPO-34 is preferred. [0034] The preferred OTP conversion catalyst is preferably incorporated into porous solid particles in which the catalyst is present in an amount effective to promote the desired OTP reactions, hi one aspect, the porous solid particles comprise a catalytically effective amount of the molecular sieve catalyst and at least one matrix material, preferably selected from the group consisting of binder materials, filler materials, and mixtures thereof to provide a desired property or properties, e.g., desired catalyst dilution, mechanical strength, and the like to the solid particles. Such matrix materials are porous in nature and may or may not be effective to help promote the desired OTP conversion. Filler and binder materials include, for example, synthetic and naturally occurring substances such as metal oxides, clays, silicas, aluminas, silica-aluminas, silica-magnesias, silica-zirconias, silica-thorias, silica-berylias, silica-titanias, silica-alumina- thorias, silica-alumina-zirconias, aluminophosphates, mixtures of these and the like.
If matrix materials, e.g., binder and/or filler materials, are included in the catalyst composition, the non-zeolitic and/or zeolitic molecular sieve catalyst preferably comprise about 1 % to 99%, more preferably about 5% to about 90% and still more preferably about 5% to about 50%, by weight of the total composition.
The most preferred zeolitic dual-function catalyst for use in the OTP reaction step of the present invention is a zeolite having the structural configuration of ZSM-5, sometimes in the literature referred to as having a "pentasil-type" structure and disclosed in US 2003/0139635Al. A borosilicate zeolite having the ZSM-5 structural configuration is disclosed as a particularly preferred dual-function catalyst in US-A-4,433,188. The use of a zeolitic catalyst having the mordenite structural configuration is disclosed in GB-A-2171718.
Another particularly preferred class of dual-function catalyst for use in the present invention is the ELAPO molecular sieves. These materials catalyze both the direct conversion of oxygenates to light olefins and the interconversion of olefins to a desired product olefin as taught in US-A-4,677,243 and US-A-4,527,001. US-B-6,455,749 teaches the use of silicoalumino phosphate catalyst (SAPO) as a dual-function catalyst and specifically has a preference for SAPO-34 and teaches both the use of a SAPO-34 type of catalyst system and a ZSM-5 bound with silicalite type of catalyst system for use in the interconversion of C4 olefins to other olefins.
Best results with a non-zeolitic catalytic system are obtained when SAPO-34 is utilized as the dual-function catalyst. Best results with a zeolitic material are obtained with a ZSM-5 type of material. A particularly preferred feature of the invention uses a mixture of a zeolitic catalyst system with a non-zeolitic catalyst system. This mixed catalyst embodiment can use either a physical mixture of particles containing the zeolitic material with particles containing the non-zeolitic material or the catalyst or a formulation that mixes the two types of material into a binder in order to form particles having both ingredients present therein. In either case the preferred combination is a mixture of ZSM-5 with SAPO-34 in relative amounts such that SAPO-34 comprises 30 to 70 wt-% of the molecular sieve portion of the mixture with a value of about 45 to 55 wt-% being especially preferred.
The present invention uses moving bed technology in the OTP conversion step to enhance the selectivity of the process for propylene production. The use of moving bed technology is shown in US-A-5,157,181. Suitable moving bed reactors and regeneration systems have been widely employed commercially for example in the catalytic reforming of naphtha fractions to increase octane number and to facilitate the dehydrogenation of light paraffins to make olefins.
Moving bed reaction zones can have various configurations. The catalyst particles can be introduced to an upper section of the reaction zone and fed by gravity through the entire volume of the reaction zone, wherein the catalyst is contacted with the feedstream either in a countercurrent direction to the catalyst movement or in a concurrent direction. In a preferred aspect of the present invention the feedstream flow is countercurrent to the catalyst flow with the feedstream introduced into a lower portion of the reaction zone and withdrawn from an upper portion thereof. This preferred configuration can provide substantial advantages in OTP conversion reactions by contacting the feedstream with partially deactivated catalyst during the initial stages of the conversion when the driving force is high and more
active catalysts during the subsequent stages of the conversion when the driving force is lower.
More typically the catalyst particles are introduced into an annulus defined by concentric catalyst retaining screens that runs through the reaction zone, wherein the catalyst particles travel down through the annulus and are withdrawn from a lower section of the OTP reaction zone. The feedstream is introduced either to the upper or lower section of the reaction zone and is passed across the annulus generally in a direction transverse to the catalyst flow. The radial bed configuration can provide low pressure drop across the reaction zone, hence good flow distribution.
During the OTP conversion zone traversal, a carbonaceous material, i.e., coke, is deposited on the catalyst as it moved downward through the reaction zone. The carbonaceous deposit material reduces the number of active sites on the catalyst thereby affecting the extent of the conversion and the selectivity to propylene. Thus during the moving bed conversion process a portion of the coked catalyst is withdrawn from the OTP reaction zone and regenerated in a regeneration zone to remove at least a portion of the carbonaceous material.
Preferably the carbonaceous material is removed from the catalyst by oxidative regeneration wherein the catalyst which is withdrawn from the reactor is contacted with an oxygen-containing gas at sufficient temperature and oxygen concentration to allow the desired amount of the carbonaceous materials to be removed from the catalyst. Depending upon the particular catalyst and conversion it can be desirable to substantially remove the carbonaceous material, e.g., to less than 1 wt-% and more preferably less than 0.5 wt-%. In some cases it is advantageous to only partially regenerate the catalyst, e.g., to remove from about 30 to 80 wt-% of the carbonaceous material. Preferably, the regenerated catalyst will contain about 0 to 20 wt-% and more preferably from about 0 to 10 wt-% carbonaceous material. Preferably in most instances when relatively large concentrations of carbonaceous material (i.e. coke) are present on the catalyst, that is, greater than about 1 wt-% carbonaceous material on the catalyst, carbon burn-off occurs with an oxygen-containing gas stream containing a relatively low concentration of oxygen. Preferably the oxygen content of oxygen containing gas is regulated through the use of inert gases or the recycling of flue gas materials to maintain an initial oxygen level of from about 0.5 to 2 vol.-%. By using the small concentration of oxygen, it is possible to reasonably control the oxidation of carbonaceous materials upon the catalyst without allowing excessively high
temperatures to occur, thereby preventing the possibility of permanently hydrothermally damaging the molecular sieve catalyst. Temperatures used during regeneration should be in the range from 400° to 700° C with best results obtained at about 500°to 650° C. The regeneration zone is preferably configured as a moving bed zone similar to the moving bed configuration used in the OTP reaction zone with coked catalyst fed to an upper portion of the regeneration zone and passed by gravity feed through the regeneration zone wherein the carbonaceous material is removed and the regenerated catalyst is withdrawn from a lower section of the regeneration zone and recirculated to the OTP reaction zone.
DETAILED DESCRIPTION OF THE DRAWING
The attached drawing shows a schematic representation of a flow scheme of the present invention in which a first reaction zone, 30, contains a fixed bed of etherification catalyst and a second reaction zone comprising reactors 1, 2 and 3 contains dual-function OTP conversion catalyst in a moving bed configuration. The three moving bed reactors 1, 2 and 3, are in serial configuration both with respect to the oxygenate feed and to the flow of catalyst through the reactors. The illustrated vertical cross section of the reactors shows catalyst descending through an annulus 33 which is maintained with appropriate concentric catalyst retaining screens. All three moving bed reactors operate with the charge to the reactor flowing in countercurrent relationship to the descending stream of catalyst particles. Preferably all three moving bed reactors have the oxygenate-containing feed stream flowing from the outside of the catalyst annulus 33 into the center portion from which the effluent stream is recovered. The drawing shows the flow of feed materials, intermediate materials and product materials by solid lines and the flow of catalyst to and from the reaction zones by dotted lines. A transport medium, preferably steam, nitrogen or any other available inert diluents transports the catalyst. The preferred catalyst transport medium is steam due to its substantial presence in the OTP reaction zone.
The etherification catalyst used in the first reaction zone, zone 30, is an acidic porous alumina material used in the form of spherical particles having a diameter of about 1.6 mm (0.0625 in) comprising gamma-alumina and having a surface area of about 180 m2 g.
The catalyst utilized in the second reaction zone comprising reactors 1, 2 and 3 is selected from the dual-function catalyst previously described and is utilized in a spherical form having a effective diameter of about .5 to 5 mm with a diameter of 1.6 mm being especially preferred. The total amount of OTP catalyst is preferably divided equally among the three reactors.
The number of reactors in the second reaction zone is selected to hold the conversion conditions in the individual reactors at conditions which enhance the yield of propylene by limiting the temperature differential across the individual reactors to 60° C or less thereby avoiding a shift in the yield structure towards ethylene and simultaneously minimizing the coke formation on the catalyst which accelerates rapidly as previously explained.
Line 8 charges a methanol-rich stream containing 95 wt-% methanol in admixture with 5 wt-% water to the process. During startup of the process recycle lines 14, 17, 21 and 22 are blocked off until sufficient product material is obtained for initiation of recycle. Similarly water diluent recycle through lines 13, 28 and 29 will be blocked off and instead an outside source of either water or steam will be injected into line 13 by means not shown just prior to the interconnection with line 31. At startup an alcoholic oxygenate feedstream flows via line 8 into zone 30 into contact with acidic porous alumina catalyst contained therein at etherification conditions effective to selectively convert at least about 10 to 70% or more of the alcoholic oxygenate to the corresponding ether. The etherification conditions in zone 30 include a temperature of 200 to 375° C, an inlet pressure of about 136 to 1136 kPa (5 to 150 psig) and a WHSV of about 0.5 to 5 hr-1. These conditions will convert from about 10 to 70% or more of methanol in the feedstream to dimethyl ether and water in an exothermic reaction that heats the ether-containing effluent stream 31 to a temperature from 375 to 525°C and shifts at least 10%, and preferably 20 to 30% or more, of the exothermic heat of reaction from the second reaction zone comprising moving bed reactors 1, 2 and 3 to this etherification step, thereby preheating the feedstream and avoiding detrimental temperature excursions in the second reaction zone.
Line 31 withdraws the resulting heated ether-containing effluent stream from zone 30 via line 31 and suitable means (not shown) adjust the temperature of this stream to a range of 375 to 525° C to produce a heated mixture of Dimethyl ether, unreacted methanol and steam diluent. The Dimethyl ether synthesis reaction form steam diluent as by-product water in an amount of at least 0.5 mole per mole of
methanol converted therein. The resulting heated mixture flows via line 31 to the intersection with line 13 that admixes an appropriate amount of diluent in the form of water or steam to provide an oxygenate to diluent ratio of about 0.5 :1 to about 5 :1 with a value of about 0.5:1 to 2:1 being especially preferred for startup. The resulting mixture of oxygenate feed and diluent passes through appropriate feed and effluent heat exchange and heating means (not shown) to fully vaporize the resulting stream and provide a charge stream for reactor 1 that enters the reactor at a temperature of about 375[deg.] to 525° C and a total pressure of about 136 to 1136 kPa (5 to 35 psig). Reactors 1, 2, and 3 contain sufficient dual- function catalyst to provide a Weight Hourly Space Velocity (WHSV) of about 0.5 to 5 hr-1 during startup with the effective WHSV rising to a level of about 1 to 10 hr-1 once recycle of olefinic hydrocarbon material commences. Line 9 withdraws effluent material from reactor 1, reduces its temperature to a value close to the temperature of the charge to reactor 1 via one or more cooling means not shown, and charges the resulting cooled effluent stream to reactor 2 wherein it contacts a dual-function catalyst to convert additional amounts of oxygenate material to propylene with production of an effluent stream withdrawn via line 10. Appropriate means (not shown) cool the effluent in line 10 to a temperature close to the inlet temperature to reactor 1 which passes into reactor 3 and contacts an additional amount of dual-function catalyst under conditions which result in further conversion of unreacted oxygenate to propylene and various other by-products. Maintaining similar temperature differentials across the reactor of the second reaction zone minimizes coking of the catalyst in reactors 2 and 3 to the same degree as in reactor 1.
Line 11 withdraws the effluent stream from reactor 3, an appropriate cooling step liquefies a substantial amount of the water contained therein via one or more cooling means such as a feed/effluent exchanger (not shown), and passes the effluent to a three-phase separator, zone 5, in which forms a hydrocarbon vapor phase, a liquid hydrocarbon phase and a water phase containing significant amounts of any unreacted oxygenate that escapes from reactor 3. By maintaining the activity of the dual-function catalyst at essentially start-of-cycle conditions this invention expects to pass a minimal amount of unreacted oxygenate into zone 5 and to achieve an overall conversion of oxygenate feed achieving through 1, 2 and 3 of 97% or greater during the entire cycle.
Line 13 withdraws the aqueous phase from separator 5. A drag stream taken by line 23 dispose of surplus water. Line 13 recycles the resulting net water stream into admixture with the oxygenate feed
via line 31. Lines 28 and 29 may add additional quantities of water for cooling purposes to reactors 2 and 3. The initiation of water recycle terminates the startup provisions for injection of diluent water.
A line 12 withdraws a vapor phase from separator 5 and charges it to fractionation column 6 which acts as a deethanizer and operates to produce an ethylene-rich overhead fraction 14 that also contains minor amounts of ethane and some methane and a bottom fraction 15 which essentially comprises the C3+ portion of the material charged to column 6. A line 18 withdraws a drag stream from the overhead stream 14 to control the buildup of C1 and C2 paraffins in the ethylene recycle loop. The ethylene recycle loop may require additional treatment of the ethylene rich overhead stream to remove methane from this stream and prevent the buildup of a significant methane concentration. Any suitable means may be used and include a demethanizer column, a methane adsorption zone, and a methane-selective membrane zone. Line 18 may take 1 to 15 vol-% of the overhead stream flowing through line 14 and more typically will comprise about 1 to 10 vol-% of this stream. Line 14 charges the remainder of the overhead stream to reactor 1 as an ethylene-rich recycle stream via line 31. The present invention may apportion the ethylene-rich recycle stream flowing in line 14 between the three reactor zones but the drawing depicts the preferred flow arrangement that charges the entire amount of this ethylene to reactor 1 and results in superior ethylene conversion when by its exposure to the maximum amount of available catalyst as it flows through the three reactor zones.
Line 15 charges the C3- material of the bottom stream to depropanizer column 7 at fractionation conditions that produce a propylene-rich overhead stream 16 as the principal product stream of the present invention and which contains minor amounts of by-product propane materials. Line 17 withdraws the bottom stream of C4+ olefin-rich material which primarily comprises C4, C5 and C6 olefinic material along with very minor amounts of butane and pentane from column 7. A line 19 takes a drag stream in an amount of about 1 to 15 vol-% and preferably about 1 to 3 vol-% to control buildup of these paraffinic materials in this C4+ olefin recycle loop. Line 17 together with lines 22 and 23 pass the remainder of the bottom material to reactors 1, 2 and 3 to provide C4+ olefin reactants unblocking of lines21 and 22 occurs once recycle operations commence. Preferably the major portion of the bottom material from column 7 passes to reactor 1 and minor portions pass to reactors 2 and 3 with preferably 60 to 80 vol-% of the total amount of this C44 olefin rich recycle stream flowing to reactor 1 via lines 17 and equal portions of about 10 to 20 vol-% being simultaneously entering reactors 2 and 3 via lines
17, 21 and 9 for reactor 2 and via lines 17, 21, 22 and 10 for reactor 3.
Once the three recycle streams flowing through lines 13, 14 and 17 are up and running then the startup mode of the instant invention is terminated and the full recycle mode of operation commences with the dual-function catalyst located in reactors 1, 2 and 3 functioning not only as oxygenate conversion catalyst but also as olefin interconversion catalyst. In the full recycle mode an alcoholic oxygenate rich feedstream enters the process via line 8 and is converted at least in part to the corresponding ether in zone 30. The resulting ether-containing mixture flowing via line 31 admixes with a first ethlylene rich recycle stream carried in line 14 and then with a water diluent in line 13. The resulting admixture passes to the intersection of line 31 where mixture with an additional quantity of C4+ olefinic material from line 17 forms the charge to reactor 1. This charge mixture flows into reactor 1 in a manner previously described after suitable heating to the appropriate inlet temperature.
After passage through the bed of catalyst maintained in reactor 1 line withdraws the resulting effluent. Lines 28 adds an additional quantity of relatively cool water diluent along and line 21 adds an additional quantity of the relatively cool C4+olefin-rich recycle stream. After suitable cooling to achieve the prescribed inlet temperature for reactor 2 set forth the resulting mixture enters reactor 2 wherein it makes a passage through the bed of catalyst to produce an effluent stream. Line 10 withdraws the effluent and flows into admixture with an additional quantum of relatively cool water diluent which from line 29 and an additional quantum of relatively cool C4+ olefin-rich material from line 22. After cooling of the resulting mixture to the prescribed inlet temperature line 10 delivers the mixture to reactor 3 for contact with the catalyst therein and production of an effluent stream which after appropriate quench and cooling flows through line 11 to three-phase separating zone 5.
The amount of dual-function catalyst utilized in reactors 1, 2 and 3 can vary. Reactors 2 and 3 may contain larger amounts of catalyst to make up for the slight amount of deactivation that occurs in reactor 1 when the catalyst flows to reactor 2 via line 25 and in reactor 2 when the catalyst flows from reactor 2 to reactor 3 via line 26. Preferably reactors 1, 2, and 3 run with essentially equal amounts of catalyst in the three zones or with a division which of approximately 25 to 30% of the total catalyst in each of reactors 1 and 2 and the remaining 40 to 50 vol-% in reactor 3.
The OTP process starts catalyst circulation after the reactors are lined out at operating conditions. Catalyst circulates between the reactors via lines 25 and 26 and to regenerator zone 4 via line 27. Regeneration zone 4 removes coke deposits from the coke- containing catalyst charged thereto via line 27 using a low severity oxidative procedure as previously explained. The regeneration zone 4 also receives by means not shown an oxygen- containing gas stream containing about 0.5 to 2 vol-% oxygen which and supplied in an amount sufficient to support combustion of a major portion of the coke charged to this zone. Line 24 completes a catalyst circulation circuit defined by lines 25, 26, 27 and 24 by recirculating catalyst to reactor 1 via. A flue gas stream is withdrawn from zone 4 via a line not shown.
The flow of catalyst around this catalyst circulation circuit is selected to provide a catalyst on-stream cycle time of 300 hours or less to maintain catalyst activity, oxygenate conversion and propylene selectivity at or near start of cycle conditions. Accordingly the catalyst particles flow around this circuit such that a catalyst particle's residence time in reactors 1, 2 and 3 is not more than 300 hours before it returns to zone 4 for regeneration.
Line 32 withdraws the hydrocarbon phase from three-phase separator zone 5. This material generally boils in the gasoline range and can comprise a gasoline product stream of the present invention which may require further treatment due to the high content of olefinic material that is present therein. Preferably the liquid hydrocarbon phase of line 32 passes to an additional fractionation step (not shown) to recover a C4 to Cg olefin-rich overhead stream which can return to reactors 1, 2 and 3 to provide additional conversion of heavy olefins to propylene. Recycling of any such C4 to C6 olefin-rich stream may occur by directly charging it to column 7 or admixing it with the stream flowing in line 17.

We claim:
1. A continuous process for selective conversion of an alcoholic oxygenate feed to propylene comprising the steps of:
a) contacting the feed with an acidic etherification catalyst in a first reaction zone at etherification conditions effective to form an ether-containing effluent stream and to shift at least 10% of the exothermic heat of reaction liberated when the feed is converted to propylene from the subsequent propylene synthesis step to this etherification step, thereby heating the ether-containing effluent stream to a temperature of 250 to 450°C and producing by-product water in an amount of at least 0.5 mole per mole of alcoholic oxygenate converted;
b) adjusting the temperature of the resulting ether-containing effluent stream to a range of 375 to 525°C and adding diluent thereto to produce a heated mixture of ether, unreacted alcoholic oxygenate and diluent;
c) reacting the resulting heated mixture with dual-function catalyst particles containing a molecular sieve, having the ability to convert the oxygenates contained therein to C3 olefin and to interconvert C2 and C4+ olefins to C3 olefin, in a second reaction zone containing at least one moving bed reactor wherein the reaction zone is operated at oxygenate conversion conditions effective to convert the oxygenates contained in the mixture to propylene and at a catalyst circulation rate through the second reaction zone selected to result in a catalyst on-stream cycle time of 300 hours or less to produce a propylene-containing effluent stream containing major amounts of a C3 olefin product and water, lesser amounts of a C2 olefin, C4+ olefins and C1 to C4+ saturated hydrocarbons and minor amounts of unreacted oxygenate, by-product oxygenates and aromatic hydrocarbons;

d) passing the propylene-containing effluent stream to a separation zone and therein cooling and separating this effluent stream into a vaporous fraction rich in C3 olefin, a water fraction containing unreacted oxygenates and by-product oxygenates and a liquid hydrocarbon fraction containing heavier olefins, heavier saturated hydrocarbons and minor amounts of aromatic hydrocarbons;
e) recycling at least a portion of the water fraction recovered in step d) to step b) to provide a portion of the diluent added therein;

f) separating the vaporous fraction into a C2 olefin-rich fraction, a C3 olefin-rich product fraction and C4+ olefin-rich fraction;
g) recycling a portion of the C2 olefin-rich fraction or of the C4+ olefin-rich fraction or of a mixture of these fractions to step c); and
h) withdrawing coke-containing dual-function catalyst particles from the second reaction zone, oxidatively regenerating the withdrawn catalyst particles in a regeneration zone and returning a stream of regenerated catalyst particles to the second reaction zone.
2. The process as claimed in claim 1 wherein the alcoholic oxygenate is an aliphatic alcohol containing 1 to 6 carbon atoms.
3. The process as claimed in claim 2 wherein the aliphatic alcohol is methanol the acidic etherification catalyst is an alumina catalyst and first reaction zone produces a -dimethyl ether (DME) containing effluent.
4. The process as claimed in claims 1,2 or 3 wherein the etherification conditions include a temperature of 200 to 375°C, an inlet pressure of about 136 to 1136 kPa (5 to 150 psig) and a weight hour space velocity (WHSV) of 0.1 to 10 hr'l.
5. The process as claimed in claims 1,2, or 3 wherein the dual function catalyst contains a zeolitic molecular sieve and or an ELAPO molecular sieve.
6. The process as claimed in claim 5 wherein the zeolitic molecular sieve has a structure corresponding to ZSM-5 and the ELAPO molecular sieve is a SAPO material having a structure corresponding to SAPO-34.
7. The process as claimed hi claims 1,_2, or 3 wherein the second reaction zone contains at least three moving bed reactors, the moving bed reactors are connected in a serial flow configuration with respect to the stream of catalyst particles that passes there through, and the moving bed reactors are connected in a serial flow configuration with respect to the flow of the heated mixture fed thereto.

8. The process as claimed in claim 1,_2, or 3 wherein the liquid hydrocarbon fraction separated in step d) is further separated into a C4 to C6 olefin-rich fraction and a naphtha product fraction and least a portion of the C4 to C6 olefin-rich fraction is recycled to step c).
9. The process as claimed in claim 8 wherein a portion of the C2 olefin-rich fraction, the C4+ olefin-rich fraction and the C4 to C6 olefin-rich fraction are recycled to step c).
10. The process as claimed in claim 1,_2, or 3 wherein the dual-function catalyst particles are
regenerated in the regeneration zone using an oxygen-containing stream under conditions selected to
produce a regenerated catalyst containing less than 0.5 wt-% carbonaceous material.
11. A continuous process for selective conversion of an alcoholic oxygenate feed to propylene, substantially as hereinbefore described with reference to the accompanying drawings.

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Patent Number

271420

Indian Patent Application Number

2249/DELNP/2007

PG Journal Number

09/2016

Publication Date

26-Feb-2016

Grant Date

19-Feb-2016

Date of Filing

22-Mar-2007

Name of Patentee

UOP LLC

Applicant Address

25 EAST ALGONQUIN ROAD, P.O.BOX 5017, DES PLAINES, ILLINOIS 60017-5017, USA

Inventors:

#	Inventor's Name	Inventor's Address
1	KALNES, TOM NELSON	UOP LLC, 25 EAST ALGONQUIN ROAD, P.O.BOX 5017, DES PLAINES, ILLINOIS 60017-5017, USA
2	WEI, DANIEL HUE	UOP LLC, 25 EAST ALGONQUIN ROAD, P.O.BOX 5017, DES PLAINES, ILLINOIS 60017-5017, USA
3	GLOVER, BRYAN KENT	UOP LLC, 25 EAST ALGONQUIN ROAD, P.O.BOX 5017, DES PLAINES, ILLINOIS 60017-5017, USA

PCT International Classification Number

C07C 1/00

PCT International Application Number

PCT/US2005/027409

PCT International Filing date

2005-08-01

PCT Conventions:

#	PCT Application Number	Date of Convention	Priority Country
1	10/943,833	2004-09-16	U.S.A.