PROCESS FOR CONVERTING OXYGENATES INTO OLEFINS

Abstract: Oxygenates are converted into lower olefins in a bed of zeolite catalyst maintained in a turbulent fluidization regime at elevated temperature.

Description: This invention relates to a process for converting oxygenated aliphatic compounds, such as methanol and dimethyl ether, into lower olefins. In particular, it provides a continuous process for producing an olefinic product rich in C₃⁺ alkenes, while minimizing byproduct aromatics and ethene. In view of the availability and low cost of synthetic methanol and its corresponding dehydration product dimethylether (DME), primary emphasis will be placed on these oxygenates in the following description of the methanol-to-olefin (MTO) process. Various zeolitic catalysts are useful for converting methanol and other lower aliphatic alcohols or corresponding ethers into olefins. Recent interest has been directed to a catalytic process for converting methanol over ZSM-5 zeolite catalysts and related to valuable hydrocarbons rich in ethene and C₃⁺ alkenes. Various processes are described in U.S. Patents 3,894,107; 3,928,483; 4,025,575; 4,252,479; 4,025,572; and in copending U.S. Patent Application Ser. No. 388,768. Significance of the methanol-to-olefins (MTO) type processes, especially for producing ethene, is discussed in Hydrocarbon Processing , November 1982, pp. 117-120. It is generally known that MTO processes can be optimized to produce a major fraction of C₂-C₄ olefins; however, a significant C₅⁺ byproduct may be coproduced, including polymethylbenzenes such as durene, as described in U.S. Patent 4,025,576. Prior process proposals have included a separation section to recover ethene and other gases from byproduct water and C₅⁺ hydrocarbon liquids. Fluidized bed MTO processes provide an important economic unit in the synthesis route for producing various chemicals, fuels and other hydrocarbon products from coal and/or natural gas. As described in copending U.S. Patent Applications Ser. Nos. 548,377 and 598,955, methanol, DME or the like may be converted to liquid fuels, particularly distillate, in a multi-stage continuous process, with integration between the major process units. The initial stage MTO process hydrocarbon olefins effluent can be fed to a catalytic oligomerization unit. Although the C₃+ mono-olefins are readily converted by ZSM-5 type acid zeolites into olefinic oligomers, ethene tends to be less reactive at moderate temperatures, and may be removed from the oligomerization system. Thus, minimizing ethene production and increasing C₃+ olefins in the MTO product stream is advantageous, and this is a main object of the present invention. A process has been found for continuous conversion of oxygenated feedstock into hydrocarbon products consisting predominately of lower olefins wherein the feedstock is contacted at elevated temperature with a turbulent fluidized bed of zeolite catalyst under conversion conditions. The improvement comprises maintaining the fluidized catalyst bed in a vertical reactor column having a turbulent reaction zone by passing feedstock vapor upwardly through the reaction zone at a velocity greater than dense bed transition velocity to a turbulent regime and less than transport velocity for the average catalyst particle. A portion of coked catalyst is withdrawn from the reaction zone; by oxidatively regenerating the withdrawn catalyst and returning regenerated catalyst to the reaction zone at a rate to control catalyst activity, propane:propene molar ratio in the hydrocarbon product is maintained at about 0.04:1 to 0.1:1 under conditions of substantially complete feedstock conversion, and C₃+ olefin production is optimized. Catalyst versatility permits the same zeolite to be used in both methanol dehydration and olefin formation. While it is possible to employ substantially different catalysts in plural stages, it is advantageous to employ a standard ZSM-5 having a silica alumina molar ratio of 70:1 or greater in a once-through fluidized bed unit to convert 98 to 100 percent, preferably at least 99.5 percent, of the feedstock oxygenate. The preferred MTO catalysts include the crystalline aluminosilicate zeolites having a silica to alumina ratio of at least 12 and a constraint index of 1 to 12. Representative of the ZSM-5 type zeolites are ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38. ZSM-5 is described in U.S. Patent 3,702,886 and U.S. Patent Re. 29,948; ZSM-11 is described in U.S. Patent 3,709,979. U.S. Patent 3,832,449 describes ZSM-12; U.S. Patent 4,076,842 describes ZSM-23; U.S. Patent 4,016,245 describes ZSM-35 and U.S. Patent 4,046,839 describes ZSM-38. These medium pore, shape-selective zeolites are sometimes known as porotectosilicates or "Pentasil" zeolites. Other zeolites, catalysts and processes suitable for converting methanol/DME into lower olefins are described in U.S. Patents 4,393,265, and 4,387,263 and European Patent Application 0081683. Zeolite ZSM-34 and ZSM-45 are known MTO catalysts, as described in U.S. Patent 4,086,186 and European Patent Application No. 83305747.4, respectively. In addition to the preferred alumino-silicates, borosilicate, ferrosilicate and "silicalite" materials may be employed. ZSM-5 type pentasil zeolites are particularly useful in the MTO process because of their regenerability, long life and stability under the extreme conditions of MTO operations. Usually the zeolite crystals have a size from 0.02 to 2 microns or more, with large crystals greater than 1 micron being preferred. In order to obtain the desired particle size for fluidization in the turbulent regime, the zeolite catalyst crystals are bound with a suitable inorganic oxide, such as silica or alumina, to provide a zeolite concentration of about 5 to 90 weight %. In a preferred aspect of the process, a 40% ZSM-5 catalyst calcined wih 60% Al₂O₃ is suitably employed. Particle size distribution can be a significant factor in achieving overall homogeneity in the fluidization. It is desired to operate the process with particles that will mix well throughout the bed. Large particles having a particle size greater than 250 microns should be avoided, and it is advantageous to employ a particle size range consisting essentially of 1 to 150 microns. Average particle size is usually about 20 to 100 microns, preferably 40 to 80 microns. Particle distribution may be enhanced by having a mixture of larger and smaller particles within the operative range, and it is particularly desirable to have a significant amount of fines. Close control of distribution can be maintained to keep about 10 to 25 wt % of the total catalyst in the reaction zone in the size range less than 40 microns. This class of fluidizable particles is classified as Geldart Group A. Accordingly, the fluidization regime is controlled to assure operation between the transition velocity and transport velocity. Fluidization conditions are substantially different from those found in non-turbulent dense beds (for example, U.S. Patent 4,423,274) or transport beds. The process is described below in greater detail by way of example only with reference to the accompanying drawings, in which Figure 1 is a schematic view of a fluidized bed reactor system; Figure 2 is a graphic plot of typical catalyst distribution in a vertical reactor; Figure 3 is a yield plot for the lower olefins vs. propane:propene ratio; Figure 4 is a PONA distribution diagram for hydrocarbon product relating to Figure 3; and Figure 5 depicts the effect of reaction temperature on C₂-C₅ olefin yield vs. propane:propene ratio. Referring to Fig. 1 of the drawings, a reactor vessel 2 is shown provided with heat exchange tubes 4. There may be several separate heat exchange, steam generating tube bundles so that temperature control can be separately exercised over different portions of the fluid catalyst bed. The bottoms of the tubes are spaced above a feed distributor grid 8 sufficiently to be free of jet action by the charged feed through the small diameter holes in the grid. Provision is made for withdrawing catalyst from above the grid 8 through a conduit 10 provided with a flow control valve 12 for passage to catalyst regeneration. Provision is also made for passing the partially regenerated catalyst to the reactor fluid bed of catalyst by a conduit 14 provided with a flow control valve 16. The regenerated catalyst is charged to the catalyst bed beneath the upper interface and sufficiently below to achieve good mixing in the fluid bed. Since the amount of regenerated catalyst passed to the reactor is small by comparison, the temperature of the regenerated catalyst does not upset the temperature constraints of the reactor operations in a significant amount. The methanol feed with or without diluent may be charged through one or more openings 20 and 22 in a bottom extended portion of the reactor. The methanol in liquid condition can be sprayed by suitable means into the bed above the grid. The charged methanol feed in vaporous condition enters the vessel by inlet means 20 and 22 in open communication with chamber 24 beneath grid 8. The charged methanol passes through reactant distributor grid 8 and into the bed of catalyst at a velocity sufficient to form a generally upwardly flowing suspension of reactant and reaction product with the catalyst particles. A plurality of sequentially connected cyclone separators 30, 32 and 34 provided with diplegs 36, 38 and 40 are positioned in an upper portion of the reactor vessel comprising dispersed catalyst phase 28. The product effluent of methanol conversion, separated from catalyst particles in the cyclone separating system, then passes to a plenum chamber 42 before withdrawal through one or more openings 44 and 46. The product effluent recovered by openings 44 and 46 is cooled and separated (in means not shown) to recover liquid hydrocarbons, gaseous material and water comprising some catalyst fines. Since conversion of the methanol is at least 95% and preferably at least 99.5%, the water phase with unreacted methanol need not be processed to recover unreacted methanol. The recovered hydrocarbon product comprising olefins and/or aromatics, paraffins and naphthenes is thereafter processed as required to provide a desired gasoline or higher boiling product. Under optimized MTO conditions the turbulent bed has a superficial vapor velocity of about 0.3 to 2 meters per second. At higher velocities, entrainment of fine particles may become excessive and beyond 1Om/s the entire bed may be transported out of the reaction zone. At lower velocities, the formation of large bubbles can be detrimental to conversion. Even fine particles cannot be maintained effectively in a turbulent bed below about 0.1 m/s. A convenient measure of turbulent fluidization is the bed density. A typical turbulent bed has an operating density of about 100 to 300 kg/m³, measured at the bottom of the reaction zone, becoming less dense toward the top of the reaction zone due to pressure drop and particle size differentiation. This density is generally between the catalyst concentration employed in dense beds and the dispersed transport systems. Pressure differential between two vertically spaced points in the reactor column can be measured to obtain the average bed density at such portion of the reaction zone. For instance, in a fluidized bed system employing ZSM-5-containing particles having a clean apparent density of 1.06 g/cc and packed density of 0.85 g/cc, an average fluidized bed density of about 190 to 200 kg/m³ is satisfactory. By virtue of the turbulence experienced in the turbulent regime, gas-solid contact in the catalytic reactor is improved, providing substantially complete conversion and enhanced selectivity. One main feature of this concept is the inherent control of bubble size and characteristic bubble lifetime. Bubbles of the gaseous reaction mixture are small, random and short-lived, thus resulting in good contact between the gaseous reactants and the solid catalyst particles. A significant difference between the process of this invention and the oxygenate conversion processes of the prior art is that operation in the turbulent fluidization regime is optimized to produce C₃⁺ lower olefins. Another advantage of operating in such a mode is the control of bubble size and life span, thus avoiding large scale gas by-passing in the reactor. The process of the invention does not rely on internal baffles in the reactor for the purpose of bubble size control such as the baffles which are employed in the prior art dense bed processes discussed above. Fluidization is a gas-solid contacting process in which a bed of finely divided solid particles is lifted and agitated by a rising stream of gas. Fluidization occurs in a bed of particulates when an upward flow of fluid through the interstices of the bed of particles attains a pressure differential and frictional resistance increment exceeding particulate weight. At this point of minimum fluidization, an increase in the fluid rate lifts or supports the particles. This condition is referred to as incipient bouyancy, since at this fluid rate, the particles are still so close as to have essentially no mobility. Where it is desired to create bed homogeneity, the fluid velocity can be increased to the point of blowing bubbles or voids into the bed which is then violently mixed as the bubbles rise. The increased fluid velocity at which bubbles first form is called the incipient-bubbling velocity. The differences between a bubbling dense bed and turbulent bed can be visualized by referring to Figure 2, which depicts a typical bed density profile for a low velocity dense bed (broken line), having a fairly sharp break at a vertical point in the reactor. By contrast the turbulent bed (solid line) density ( ) decreases at a more even rate with upward bed height (h). In the past a wide variety of materials have been tested to study fluidization effects. However, the conclusions drawn from data on the fluidization of one particulate matter were not found to be applicable to other powders which had quite different particle sizes and densities. Geldart [ 7 Powder Technology 285 (1973)] reports that the behavior of solids fluidized by gases falls into four clearly recognizable groups, characterized by the difference in density between the particular particles and the fluidizing fluid and the mean size of the particles. Powders in Geldart's group A exhibit a dense phase expansion after minimum fluidization and prior to the commencement of bubbling, group B powders bubble at the minimum fluidization velocity; the powders in group C are difficult to fluidize at all; and those in group D can form stable spouted beds. Particles useful in the process of the invention as well as fluid catalytic cracking catalysts are typical group A powders, usually having a small mean size and/or a low apparent particle density (e.g. less than about 1.4 g/cm³). These powders expand considerably before bubbling commences. As the superficial gas velocity is increased in the dense bed, eventually slugging conditions occur and with a further increase in the superficial gas velocity the slug flow breaks down into a turbulent regime. The transition velocity at which this turbulent regime occurs appears to decrease with particle size. The turbulent regime extends from the transition velocity to the so-called transport velocity. (Yerushalmi et al, Further Studies of the Regimes of Fluidization , 24 Powder Tech. 187-205 (1979). As the transport velocity is approached, there is a sharp increase in the rate of particle carryover, and in the absence of solid recycle, the bed could empty quickly. The transition from bubbling dense bed to turbulent fluidization involves a gradual breakdown of large bubbles into small bubbles and voids, providing a progressive change toward a condition of greater homogeneity. Solid mixing is vigorous, the interaction of the two phases is strong, and the contact between gas and solid is highly efficient. Since the fluidized bed of zeolite catalyst of the invention is operated in the turbulent regime, the gaseous reaction mixture is present as small dispersed bubbles whose existence is extremely short. The bubbles travel only a short distance as they constantly break-up, coalesce and reform. Therefore when practicing the process it is unnecessary to provide baffles in the reactor vessel as is often found necesary to control the large bubbles encountered in bubbling dense fluidized beds, such as those employed in the prior art processes. Several useful parameters contribute to fluidization in the turbulent regime in accordance with the process of the invention. When employing a ZSM-5 type zeolite catalyst in fine powder form such a catalyst should comprise the zeolite suitably bound or impregnated on a suitable support with a solid density (weight of a representative individual particle divided by its apparent "outside" volume) in the range from 0.6 to 2 g/cc, preferably 0.9 to 1.6 g/cc. The catalyst particles can be in a wide range of particle sizes up to about 250 microns, with an average particle size between about 20 and 100 microns, preferably in the range of 10 to 150 microns and with the average particle size between 40 and 80 microns. When these solid particles are placed in a fluidized bed where the superficial fluid velocity is 0.3 to 2 M/S, MTO operation in the turbulent regime is obtained. The velocity specified here is for an operation at a total reactor pressure of about 100 to 300 kPa. At higher pressures, a lower gas velocity may be employed to ensure operation in the turbulent fluidization regime. The reactor can assume any technically feasible configuration, but several important criteria should be considered. The bed of catalyst in the reactor can be at least about 5 to 20 m in height, preferably about 9 m. Fine particles may be included in the bed, especially due to attrition, and the fines may be entrained in the product gas stream. A typical turbulent bed may have a catalyst carryover rate up to about 1.5 times the reaction zone inventory per hour. If the fraction of fines becomes large, a portion of the carryover can be removed from the system and replaced by larger particles. It is preferred to have a fine particle separator, such as a cyclone or filter, disposed within the reactor shell to recover catalyst carryover and return this fraction continuously to the bottom of the reaction zone for recirculation at a rate of about one catalyst inventory per hour. The following Examples illustrate the invention. A series of continuous methanol conversion runs are conducted in a vertical reactor vessel under turbulent fluidized bed conditions using a catalyst comprising ZSM-5 (alpha = 15) having about 3 to 36 wt. % coke deposited by prior conversion processes. The reactor is a vertical column having a turbulent reaction zone about 9 m in height and containing catalyst with an average particle size of 75 to 80 microns. The fines less than 40 microns can vary from about 15 to 25 weight %, and less than 20 microns from 3 to 7 %. The results are tabulated below: TABLE I The average coking rate is about 1.4 weight %/day without regeneration. During these runs the coke content of the catalyst increases from 3 weight % initially to 15, 17.7 and 35.6 %, at the end of runs. Continuous or periodic withdrawal of a coked catalyst fraction is employed to maintain the desired activity, as measured by the research index (RI), expressed as the propane:propene ratio. While nitrogen gas is used to demonstrate the effect of a diluent, other gases may be employed to decrease the oxygenate partial pressure, such as steam, hydrogen or light hydrocarbon recycle gas, particularly C₁-C₄ alkanes. Methanol breakthrough is observed at high space velocities and/or low RI values. While 98-99.5 % conversion is obtainable in the RI range of 0.04 to 0.06, complete conversion is achieved at 0.06 to 0.1 RI. The effect of catalysts activity, as measured by propane:propene ratio, is depicted in Figures 3 and 4 for a series of runs conducted in the turbulent bed of Examples 1-3 except that average bed temperature is maintained at about 482°C and methanol partial pressure is maintained at about 100 kPa by N₂ addition. Figure 3 shows hydrocarbon product yield for each of the C₂-C₅ olefins, while Figure 4 shows the total product distribution. The effect of process temperature is demonstrated in Figure 5 for another similar series of runs at 185 kPa using undiluted methanol feedstock. While C₂-C₅ olefin yield is increased by raising the temperature from 480 to 500°C, a further rise in temperature to 520°C did not improve these yields.

Claims: 1. A process for converting a feedstock comprising methanol and/or dimethylether into hydrocarbons rich in light olefins, comprising the steps of maintaining a fluidized bed of zeolite catalyst particles in a turbulent reactor bed at a temperature of at least 475°C, the catalyst having an apparent particle density of 0.9 to 1.6 g/cm³ and a size range of 1 to 150 microns, an average catalyst particle size of 20 to 100 microns and containing 10 to 25 weight percent of particles having a particle size less than 32 microns; passing hot feedstock vapor upwardly through the fluidized catalyst bed under turbulent flow conditions; maintaining turbulent fluidized bed conditions through the reactor bed between transition velocity and transport velocity at a superficial fluid velocity of 0.3 to 2 meters per second; and recovering hydrocarbon product containing a major amount of olefins and having a molar ratio of C₃+ olefins to ethene of at least 12:1. 2. A process according to claim 1 wherein the fluidized bed density is 100 to 300 kg/m³, measured at the bottom of the bed, and wherein the catalyst comprises ZSM-5 type zeolite. 3. A process according to claim 1 or claim 2, wherein the feedstock comprises crude methanol at least 98 weight percent of which is converted in the turbulent bed which has a height of at least 9 meters. 4. A process according to any one of claims 1 to 3, wherein the catalyst comprises a medium pore pentasil zeolite having an apparent alpha value of 3 to 50. 5. A process according to any one of claims 1 to 4, wherein the catalyst particles comprise 5 to 90 weight percent ZSM-5 zeolite having a crystal size of 1 to 2 microns. 6. A process according to any one of claims 1 to 5, wherein the feedstock comprises a major amount of methanol and up to 50 mole percent of a diluent gas, and wherein the partial pressure of methanol is from 35 to 240 kPa.

Inventors: AVIDAN, AMOS ANDREW, GOULD, RONALD MICHAEL, KANE, STEVEN EDWARD
Applicants: MOBIL OIL CORPORATION
Application year: 1985
Main classification: C07C1
IPC codes: C07C1