PROCESS AND REACTOR SYSTEM FOR CONVERTING METHANOL TO OLEFINIC HYDROCARBONS

Abstract: A process for converting crude aqueous methanol feedstock or the like to olefinic hydrocarbons in contact with a medium pore shape selective crystalline acid zeolite catalyst,wherein the improvement comprises contacting the aqueous methanol feedstock with a liquid propane-rich hydrocarbon extractant under liquid extraction conditions, recovering an aqueous phase containing the major amount of water introduced with the feedstock, recovering an organic extract phase comprising the hydrocarbon extractant and a portion of methanol introduced in the feedstock, and converting the extracted methanol at elevated temperature under catalytic reaction conditions to produce predominantly olefinic hydrocarbons.

Description: This invention relates to techniques for converting an oxygenated aliphatic compound, such as methanol, to lower olefins. In particular, it provides a continuous process for producing an olefinic product rich in C₂-C₅ alkenes. In view of the availability and low cost of synthetic methanol (MeOH), primary emphasis is placed on this feedstock material in following description of the methanol-to-olefin (MTO) process. Processes for converting lower oxygenates such as methanol to hydrocarbons are known and have become of great interest in recent times because they offer an attractive way of producing liquid hydrocarbon fuels, especially gasoline, from sources which are not of liquid petroleum origin. In particular, they provide a way by which methanol can be converted to a major amount of C₂-C₅ olefins and a minor amount of gasoline boiling range products in good yields. The methanol, in turn, may be readily obtained from coal by gasification to synthesis gas and conversion of the synthesis gas to methanol by well-established industrial processes. As an alternative, the methanol may be obtained from natural gas by other conventional processes, such as steam reforming or partial oxidation to make the intermediate syngas. Crude methanol from such processes usually contain a significant amount of water, usually in the range of 4 to 20 wt.%; however, the present invention is useful for removing water in lesser amounts or greater. Various zeolitic catalysts are useful for converting methanol and other lower aliphatic alcohols or corresponding ethers to olefins. Recent interest has been directed to a catalytic process for converting methanol over ZSM-5 and related catalysts to valuable hydrocarbons rich in ethene and C₃⁺ alkenes. Various processes are described in U.S. Patents 3,894,107 (Butter et al); 3,928,483; 4,025,575; 4,052,479 (Chang et al); 4,025,572 (Lago); 4,328,384 (Daviduk et al); and 4,547,616 (Avidan et al). 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. Prior process techniques for increasing lower olefin selectivity have provided for controlled deposition of coke byproduct on the catalyst surface. Methanol may be first subjected to a dehydrating step, using a catalyst such as gamma-alumina, to form an equilibrium mixture of methanol, dimethyl ether (DME) and water. This mixture is then passed at elevated temperature and pressure over a catalyst such as ZSM-5 zeolite for conversion to the hydrocarbon products. Water may be removed from the methanol dehydration products prior to further conversion to hydrocarbons and the methanol can be recycled to the dehydration step, as described in U.S. Patent No. 4,035,430. Removal of the water is desirable because the catalyst may tend to become deactivated by the presence of excess water vapor at the reaction temperatures employed; but this step is not essential. Typically the crude methanol feestock employed in MTO processes contains about 4 to 20 wt.% water as the principal impurity. Excessive water not only contributes to catalyst deactivation, but also requires larger volume equipment to handle the increased throughput. Various proposals have been put forth for reducing the water content of crude methanol, for instance the distillation system described by Mao et al in US-A-4665249. The present invention provides a continuous technique for converting crude oxygenate feedstock, such as methanol and/or dimethylether, to predominantly olefinic hydrocarbons in a catalytic reaction zone with a crystalline acid zeolite catalyst at elevated temperature. Processes and apparatus are provided for contacting a crude oxygenate feedstock containing a minor amount of water with a liquid hydrocarbon extraction stream rich in propane and lighter hydrocarbons under extraction conditions favorable to selective extraction of the oxygenate, thereby providing an extract liquid stream rich in oxygenate and an aqueous raffinate stream lean in oxygenate. Substantially free of water, the extracted oxygenate is passed to the reaction zone under process conditions to convert oxygenate to predominantly olefinic hydrocarbons. This is followed by cooling reaction effluent to recover aqueous liquid byproduct, gas rich in C₃⁻ hydrocarbons, and product comprising C₄⁺ hydrocarbons. Extractant is provided by condensing and recycling at least a portion of the C₃⁻ liquid phase for use as propane-rich extraction liquid. A major advantage of C₃⁻ recycle is that an increased circulation of C₃⁻ material through the reaction zone can provide reactant dilution and enhanced selectivity of C₂-C₅ olefins. In the drawings, FIG. 1 is a schematic MTO process flowsheet depicting the present invention; FIG. 2 is a flow diagram of the extraction and methanol dewatering subsystem for a preferred embodiment using C₃ aliphatic recycle; FIG. 3 is a graphic plot of methanol recovery vs. carbon number of n-alkane extractant, showing extraction temperature as parameter; FIG. 4 is a corresponding graphic plot showing water entrainment in the extract phase for C₃-C₇ alkanes; and FIG. 5 is a process flow diagram for an alternative technique for dewatering MeOH-DME mixtures. Suitable oxygenate feedstocks include C₁-C₄ aliphatic alcohols, ethers, aldehydes, ketones, and the like. Dehydration mixtures of methanol and DME may contain about 10-60% DME, for instance. Referring to Fig. 1 of the drawing, crude methanol (MeOH) feedstock is introduced via conduit 10 to a top inlet of extraction unit 12, where it is contacted under liquid extraction conditions with a countercurrent stream of hydrocarbon liquid extractant introduced at a bottom inlet via conduit 14. An aqueous raffinate phase containing a major amount of the water present in the crude feedstock and a portion of MEOH is withdrawn via conduit 16. Raffinate methanol can be recovered in distillation tower 17 and combined with the extract phase in conduit 18. The lighter organic extract phase containing hydrocarbon extraction solvent and methanol is recovered from extraction unit 12 via conduit 18, passed through heat exchanger 20, and introduced at elevated temperature and process conditions suitable for conversion of substantially all of the methanol in contact with the MTO catalyst in reactor system 30. From the reactor system 30, the effluent product stream leaves via line 32 to a heat exchanger 34 and condensate accumulator 34A. Vapor is pressurized in primary compressor 35, cooled in intermediate condenser 36, and passed as a three-phase mixture to a primary phase separator 37 for recovery of byproduct water and condensed liquid. Secondary compressor 38 raises the vapor to a suitable pressure (eg-1500-2200 kPa; 200-350 psig) for feeding to depropanizer tower 40 for recovery of C₃⁻ gas overhead rich in C₃ hydrocarbons, especially propane and C₂-C₃ olefins for recycle. Advantageously, at least a portion of the C₃ tower overhead components is condensed by cooling with exchanger 41 and passed via accumulator 42 for use as reflux and liquid extractant. Propene contained in the C₃ liquid recycle stream can be further converted to C₄⁺ hydrocarbons in reactor 30. Light gas 48 from the accumulator and optional C₃⁻ purge stream 46 may be removed from the loop. A portion of ethene-rich overhead vapor may be recycled via line 44 directly to the MTO reactor 30 for further conversion. This system configuration is particularly advantageous since it eliminates expensive cryogenic necessary for complete separation of C₂-C₃ components. The C₃⁻ purge stream 46 and ethene-rich offgas stream 48 may be utilized as fuel gas, or olefinic components can be further upgraded by oligomerization in a high severity reactor. Tower 40 bottoms comprise C₄⁺ hydrocarbons, which may be further fractionated to recover isobutane or other components thereof. A recycle C₃⁻ stream including at least a portion of liquid propane recovered from the process reactor effluent is sent via conduit 14 for use as extractant liquid, along with any added propane or other makeup hydrocarbon liquid added to the process extractant stream. The aqueous raffinate stream 16 consists essentially of water, partitioned methanol and a trace of hydrocarbon. This may be recovered and fractionated to provide a methanol-rich stream to be combined with crude feestock in conduit 10. Separaton of C₃⁻ recycle from reactor effluent may be facilitated by employing an intermediate absorber unit (not shown) between the primary compressor and depropanizer tower for contacting compressed vapor with a liquid hydrocarbon sorption stream for selective sorption of C₄⁺ components. The preferred crude feedstock material is methanol containing about 4 to 17% by weight water. The extraction unit 12 is depicted as a stirred multi-stage vertical extracton column adapted for continuous operation at elevated pressure. Any suitable extraction equipment may be employed, including concurrent, cross-current or single stage contactors, wherein the liquid methanol feedstock is intimately contacted with a substantially immiscible liquid hydrocarbon solvent, which may be a mixture of C₃-C₄ aliphatic components or relatively pure propane, butane, etc. This unit operation is described in Kirk-Othmer Encyclopedia of Chemical Technology (Third Ed.), 1980, pp. 672-721. Other equipment for extraction is disclosed in US Patents 4,349,415 (DeFilipi et al) and 4,626,415 (Tabak). The methanol extraction step can be performed advantageously in a countercurrent multistage design, such as a simple packed column, rotating disk column, agitated column with baffles or mesh, or a series of single stage mixers and settlers. As an example of crude methanol dewatering technique, Fig. 2 shows an integrated extraction and distillation operation with thermal recovery, employing similar ordinal numbers for corresponding elements in Fig. 1. In this example crude methanol containing about 17 wt.% water is fed via inlet 110 to extraction unit 112 at a flow rate of 16364 moles MeOH + 6080 moles H₂O, where it is contacted with 89776 moles of C₃ (propane) hydrocarbon at a temperature of 43°C. The extract phase containing 9596 moles of MeOH, 325 moles H₂O (less than 2 wt.%), and 89560 moles C₃ hydrocarbon is passed via conduit 118, heat exchangers 121, 123 and combined with distilled methanol to provide a reactor system feedstream containing 16357, 1175 and 89776 moles, respectivlely. Raffinate stream 116, containing 6768 moles MeOH, 5755 moles H₂O, and 216 moles C₃ is heated in exchanger 117 and introduced to distillation tower 150 for further enrichment of the methanol overhead, which is condensed in exchanger 123 for reflux and reactor feed. Distillation bottoms are reboiled in exchanger 151 and a wastewater stream 155 is recovered containing about 4905 moles of water and 7 moles of MeOH. Recent developments in zeolite technology have provided a group of medium pore siliceous materials having similar pore geometry. Most prominent among these intermediate pore size zeolites is ZSM-5, which is usually synthesized with Bronsted acid active sites by incorporating a tetrahedrally coordinated metal, such as Al, Ga, or Fe, within the zeolytic framework. These medium pore zeolites are favored for acid catalysis; however, the advantages of ZSM-5 structures may be utilized by employing highly siliceous materials or crystalline metallosilicate having one or more tetrahedral species having varying degrees of acidity. ZSM-5 crystalline structure is readily recognized by its X-ray diffraction pattern, which is described in U.S. Patent No. 3,702,886 (Argauer, et al.). Catalyst versatility permits the same zeolite to be used in both methanol dehydration and olefin formation. While it is within the inventive concept 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 feedstock oxygenate to hydrocarbons. The MTO catalysts preferred for use herein include the crystalline aluminosilicate zeolites having a silica to alumina ratio of at least 12, 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, ZSM-38, and ZSM-48. ZSM-5 is disclosed and claimed in U.S. Patent No. 3,702,886 and U.S. Patent No. Re. 29,948; ZSM-11 is disclosed and claimed in U.S. Patent 3,709,979. Also, see U.S. Patent 3,832,449 for ZSM-12. Also, see U.S. Patent No. 3,832,449 for ZSM-12; U.S. Patent No. 4,076,842 for ZSM-23; U.S. Patent No. 4,016,245 for ZSM-35; U.S. Patent No. 4,046,859 for ZSM-38 and U.S. Patent No. 4,375,573 for ZSM-48. The medium pore shape selective catalysts are sometimes known as porotectosilicates or "pentasil" catalysts, and may have a pore size of about 5 to 7 Angstroms. Other catalysts and processes suitable for converting methanol/DME to lower olefins are disclosed in U.S. Patents 4,393,265 and 4,613,720 (Bonifaz), U.S. 4,387,263 (Vogt et al) and European Patent Application Publ. No. 0081683 (Marosi et al). ZSM-34 and ZSM-45 are known MTO catalysts, as disclosed in U.S. Patent 4,086,186 (Rubin et al) and EP-A-107370, respectively. In addition to the preferred aluminosilicates, silicoaluminophosphate, gallosilicate, 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 crystal size from 0.02 to 2 microns or more, with large crystals on the order of 0.1-1 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, alumina, etc. to provide a zeolite concentration of about 5 to 90 wt.%. It is advantageous to employ a particle size range of 1 to 150 microns. Average particle size is usually about 20 to 100 microns, preferably 40 to 80 microns. The catalyst in the fluidized bed reactor is maintained at an average acid cracking activity (alpha value) of about 1 to 15, preferably about 3 to 8, on a coke-free basis. The average coke content is less than 15 weight percent, preferably about 5-10 wt.% of the clean-burned catalyst. By controlling the catalytic properties of the system, the selectivity to produce C₂-C₅ olefins can be enhanced. MTO conversion in the presence of lower aliphatic diluents is described in U.S. Patents 4,025,575 (Chang et) and in EP-A-229952. In the following example H-ZSM-5, prepared according to US Patent 4,375,458 (Dwyer) is modified to deposit MgO(5.1% Mg) into the pores. Example 1 Dewatered methanol is converted over the Mg-modified H-ZSM-5 at about 350°C and 100 kPa in the presence of propane at a molar ratio of 3.1 C₃: MeOH, with methanol space velocity of about 0.4 WHSV. Under these condtions MeOH conversion is 100%, while a small amount of propane is produced. The C₂-C₄ olefins selectivity is about 61 wt.% in addition to about 29% butane and 5% C₅⁺ liquid product. Example 2 Comparison is made between a fluidized bed reactor operation with and without C₃ recycle. The data in Table 1 are directed to continuous conversion of dewatered methanol employing a particulate 25% HZSM-5 catalyst having an acid cracking value of 4 under steady state conditions, as described in US Patent 4,547,616 (Avidan et al), with recovery and recycle of ethene. Fuel gas make decreases by 4 wt.%; and isobutane product increases by about 19 wt.% in the case using 3 moles of propane diluent per mole of MeOH. Propane dilution appears to concentrate olefin production in the propene to butene range (ethene is recycled to extinction); while reducing fuel gas make and increasing isobutane production. For cases involving the use of the propane diluent, the possibility of using a fixed bed reactor vs. a fluidized bed reactor should be considered. If the propane is recycled for improved selectivity, it will also help to control the temperature gradient across a fixed bed adiabatic reactor. In a preferred embodiment, extracted methanol is passed to the reactor system with hydrocarbon extractant containing recycle. The combined feedstock and recycle is conducted at a temperature of 275-525°C, preferrably 475-500°C, and a pressure of 100-1000 kPa to the MTO catalytic reactor. Effluent from the reaction zone is passed to a heat exchange condenser system and then to a separation zone without excessive loss of pressure. In a high propane recycle ratio operation the reactor effluent remains in vapor phase after cooling, and the entire stream is compressed. At low propane recycle ratio the effluent stream may be partitioned into three streams in the separation zone. Typically, the separator is operated at conditions of about 38°C (100°F) and about 100-500 kPa (15-70 psia). The operation of the MTO unit is at low pressures and the effluent should be compressed for further processing. The increase in compression requirements and reactor size are the two process conditions most affected by the use of the propane diluent. The C₃⁻ recycle stream may comprise, on a molar basis, 25 to 90% of the reactor feedstream, preferrably 35 to 80 mole%, preferrably including propane in a molar ratio to neat methanol of about 0.5 to 4: 1. It may be desirable to further fractionate the C₃⁻ to recover and/or alkylate the isobutylene (see US 4,633,028, Owen et al) or to upgrade olefins (see US 4,579,999, Gould et al). Similarly, isobutane produced by oxygenate conversion or from other aliphatics may be separated for other use. Where a supply of outside propane is available, a major amount of the extractant can be comprised of recycled and makeup propane. Typically, 95% of recycle are C₂-C₃ aliphaitic hydrocarbons. In the alternative embodiment of Fig. 5 crude methanol is partially converted to dimethyl ether (DME) by a dehydration reaction step to provide a mixture of MeOH, DME and H₂O prior to dewatering. Crude methanol introduced via line 210 may be fed directly to the DME reactor 222 or to the extraction unit via line 214 or fed to methanol recovery tower 217 along with aqueous raffinate stream 216. MeOH-rich overhead from tower 217 is heated in exchanger 220 and contacted with gamma-alumina or other acid metal oxide dehydration catalyst in DME reactor 222 to provide an intermediate aqueous MeOH-DME oxygenate mixture, which is cooled by exchanger 224 and mixed with liquid hydrocarbon extractant (eg-C₃ aliphatics) in a stirred mixing vessel 226 to effect thorough mixing of the oxygenate and hydrocarbon phaser. This liquid mixture is separated in phase separator unit 212 to provide an aqueous raffinate 216 and oxygenate-rich extract phase 218. After heating in exchanger 219, this mixture is passed via line 244 to the MTO reactor unit 230 for conversion of the oxygenates to lower olefins. The MTO reactor effluent is processed as in Fig. 1 to recover product and recycle streams.

Claims: 1. A continuous process for converting crude methanol to olefinic hydrocarbons in a catalytic reaction zone with a crystalline acid zeolite catalyst at elevated temperature comprising the steps of: (a) contacting a crude methanol feedstock containing a minor amount of water with a liquid hydrocarbon extraction stream rich in propane and lighter hydrocarbons under extraction conditions favorable to selective extraction of the methanol, thereby providing an extract liquid stream rich in methanol and an aqueous raffinate stream lean in methanol; (b) charging the extract liquid stream rich in methanol to the catalytic reaction zone under process conditions for converting methanol to predominantly C₂-C₅ olefinic hydrocarbons; (c) cooling the reaction effluent from (b) to recover an aqueous liquid byproduct, a gas rich in C₃⁻ hydrocarbons, and a liquid product comprising C₄⁺ hydrocarbons; and (d) condensing and recycling at least a portion of the C₃ liquid phase to step (a) for use as a propane-rich extraction liquid. 2. The process of claim 1 wherein the catalyst comprises a ZSM-5 type zeolite. 3. The process of claim 1 or 2 wherein the feedstock comprises methanol and 4 to 20 wt% water, and wherein the extraction liquid comprises a major amount of propane C₃ hydrocarbons. 4. The process of claim 1, 2, or 3 wherein a major amount of C₃⁻ hydrocarborns from the reaction effluent is recycled with extracted methanol to the reaction zone as a vapor stream. 5. The process of any one of the preceding claims comprising the additional steps of: (a′) separating the residual methanol form the aqueous raffinate stream for conversion; (b′) heating substantially the entire extract liquid stream and the residual methanol from the aqueous raffinate stream to provide a reactant stream comprising hot methanol and light propane-rich hydrocarbon vapor; (b″) converting the extracted methanol and residual methanol from the raffinate in the presence of the light hydrocarbon vapor at elevated temperature under catalytic reaction conditions to produce predominantly olefinic hydrocarbons; and (c′) separating the conversion reaction effluent to recover a C₄⁺ olefin-rich product, C₃⁻ light gas, and a water byproduct. 6. The process of claim 5 wherein step (a′) is conducted in a continuous extraction unit having means for contacting the extractant and feedstock under continuous conditions. 7. The process of claim 5 or 6 wherein the catalyst consists essentially of HZSM-5 and wherein the crude feedstock contains 4 to 20 w% water, and 10 to 60 wt% dimethyl ether. 8. A continuous feedstock and methanol to olefins reactor system for converting crude oxygenated hydrocarbon feedstock to olefinic hydrocarbons comprising: extractor means (12) for contacting crude feedstock liquid containing a minor amount of water with a liquid hydrocarbon extraction stream under extraction conditions favorable to selective extraction of oxygenate, thereby providing an extract liquid stream rich in oxygenate and an aqueous raffinate stream lean in oxygenate; catalytic reactor means (30) for contacting the extract stream in a catalytic reaction zone with a crystalline acid zeolite catalyst at elevated temperature in a reaction zone under process conditions to convert a major portion of oxygenate to hydrocarbons; separation means (34, 34A) for receiving reactor effluent to recover an aqueous liquid byproduct stream, a gaseous stream rich in C₃⁻ hydrocarbons and a product stream comprising C₄⁺ hydrocarbons; and means (14) for recycling at least a portion of C₃ liquid hydrocarbons recovered from reactor effluent to said extractor means (12) for use as extraction liquid. 9. The system of claim 8 wherein the extractor means comprises a countercurrent liquid-liquid extraction column. 10. The system of claim 8 or 9 further comprising depropanizer means (40) for recovering C₄⁺ liquid product and C₃⁻ overhead components, including means for condensing at least a portion of propane from the C₃⁻ overhead for liquid recycle. 11. The system of claim 10 further comprising means (41, 42) for recycling at least a portion of light gas from the depropanizer means directly to the reaction zone. 12. The system of claims 8, 9, 10, or 11 including distillation means (17) for deparating residual oxygenate from the aqueous raffinate stream.

Inventors: HARANDI, MOHSEN NADIMI, OWEN, HARTLEY, SMYTH, SEAN CONNON
Applicants: MOBIL OIL CORPORATION
Application year: 1988
Main classification: C07C1
IPC codes: C07C11, C07C1, B01J29, C07B61, C07C67