Integrated Process for Converting Oxygenates to Liquid Hydrocarbons

This invention relates to an integrated system for converting oxygenates, such as methanol or dimethyl ether (DME), to liquid hydrocarbons. In particular it provides a continuous process for producing liquid hydrocarbon products by converting the oxygenate feedstock catalytically to an intermediate lower olefinic stream for upgrading to produce light distillate and/or gasoline products.

In order to provide an adequate supply of liquid hydrocarbons for use as synfuels or chemical feedstocks, various processes have been developed for converting coal and natural gas to gasoline and distillate. A substantial body of technology has grown to provide oxygenated intermediates, especially methanol. Large scale plants can convert methanol or similar aliphatic oxygenates to liquid fuels, especially gasoline. However, the demand for heavier hydrocarbons has led to the development of processes for making high octane gasoline and quality diesel fuel by a multi-stage technique, utilizing mainly C.sub.3.sup.+ olefins for upgrading.

Recent developments in zeolite catalysts and hydrocarbon conversion processes have created interest in utilizing olefinic feedstocks for producing C.sub.5.sup.+ gasoline, diesel fuel, etc. In addition to the basic work derived from ZSM-5 type zeolite catalysts, a number of discoveries have contributed to the development of new industrial processes.

The medium pore ZSM-5 type catalysts are particularly useful for converting methanol and other lower aliphatic alcohols or corresponding ethers to olefins. Particular interest has been directed to a catalytic process for converting low cost methanol to valuable hydrocarbons rich in ethene and C.sub.3.sup.+ alkenes. Various processes are described in U.S. Pat. Nos. 3,894,107 (Butter et al), 3,928,483 (Chang et al), 4,025,571 (Lago), 4,423,274 (Daviduk et al) and 4,433,189 (Young), incorporated herein by reference. It is generally known that the methanol-to-olefins (MTO) process can be optimized to produce a major fraction of C.sub.2 -C.sub.4 olefins. Prior process proposals have included a separation section to recover ethene and other gases from byproduct water and C.sub.5.sup.+ hydrocarbon liquids.

It has been discovered that methanol, DME or the like may be converted to liquid fuels, particularly alkylate, in a multi-stage continuous process, with integration between the major process units providing an ethene-rich recycle stream for further conversion and liquid recycle stream for interstage sorption utilizing C.sub.5.sup.+ liquid from the MTO stage as sorbent. C.sub.3.sup.+ olefins from the initial stage MTO type process hydrocarbon effluent stream, after byproduct hydrocarbon and water separation, can be fed to an alkylation stage for upgrading to heavier hydrocarbons. Ethene may be recovered by interstage separation and recycled. Advantageously, the recycled ethene is found to be reactive with methanol/DME or other oxygenates in the presence of ZSM-5 type catalysts. In effect a novel MTO and olefin upgrading system is provided wherein the ethene component may be recycled substantially to extinction.

In a preferred embodiment, the invention provides processes and apparatus for an integrated continuous technique for converting oxygenated organic feedstock to liquid alkylate hydrocarbons comprising methods and means for

(a) contacting feedstock with zeolite catalyst in a primary catalyst stage at elevated temperature to convert at least a portion of the feedstock oxygenate predominantly to an olefinic fraction rich in C.sub.2 -C.sub.4 olefins and a minor heavy hydrocarbon portion containing C.sub.5.sup.+ aliphatic and aromatic hydrocarbons;

(b) cooling and separating effluent from step (a) to provide an aqueous liquid byproduct stream, a heavy hydrocarbon liquid stream and a light hydrocarbon stream rich in C.sub.2 -C;hd 4 olefins;

(c) compressing and contacting the light hydrocarbon stream in a sorption tower with a liquid hydrocarbon sorbent stream containing a major amount of said heavy liquid stream from step (b) to selectively sorb C.sub.3.sup.+ components in a liquid sorbate stream;

(d) fractionating the sorbate stream to provide a gasoline-rich stream and a light hydrocarbon stream rich in C.sub.3 -C.sub.4 olefins;

(e) further reacting the olefinic light hydrocarbon stream from step (d) with isoparaffin in a secondary alkylation stage in the presence of liquid phase acid catalyst to convert at least a portion of C.sub.3 -C.sub.4 olefins to a heavier C.sub.7 + liquid hydrocarbon product stream comprising alkylate gasoline;

(f) recycling ethene in a gaseous stream from the sorption tower to the primary catalytic stage for further converion; and

(g) recycling a portion of gasoline from step (d) to step (c) for use as lean sorbent.

Advantageously, the primary stage catalyst comprises ZSM-5 type zeolite and ethene is recycled to the primary stage at a rate of about 1 to 20 parts ethene per 100 parts by weight of methanol equivalent in the feedstock. By fractionating gaseous effluent separated from the primary staged effluent, a recycle gas stream may be recovered containing at least 90% of ethene from the primary catalytic stage. An olefinic stream rich in C.sub.3.sup.+ olefins, especially propene and butylenes, is provided for upgrading by reaction with various isoparaffins, such as isobutane.

Other objects and features of the invention will be seen in the following description and drawings.

FIG. 1 is a process flow sheet showing the major unit operations and process streams; and

FIG. 2 is a schematic representation of an alternative inter-stage separation system.

Numerous oxygenated organic compounds may be contained in the feedstock material to be converted in the primary stage. Since methanol or its ether derivative (DME) are industrial commodities available from synthesis gas or the like, these materials are utilized in the description herein as preferred starting materials. It is understood by those skilled in the art that MTO-type processes can employ methanol, dimethylether and mixtures thereof, as well as other aliphatic alcohols, ethers, ketones, and/or aldehydes. It is known in the art to partially convert oxygenates by dehydration, as in the catalytic reaction of methanol over gamma-alumina to produce DME intermediate. Typically, an equilibrium mixture (CH.sub.3 OH+CH.sub.3 OCH.sub.3 +H.sub.2) is produced by partial dehydration. This reaction takes place in either conversion of methanol to lower olefins (MTO) or methanol to gasoline (MTG).

The process may be optimized by employing fluid bed primary stage conditions in the temperature range of about 425.degree. C. to 550.degree. C., a pressure range of about 100 to 800 kPa and weight hourly space velocity range of about 0.5 to 3.0 based on ZSM-5 equivalent catalyst and methanol equivalent in the primary stage feedstock. Suitable equipment and operating conditions are described in U.S. patent application Ser. No. 687,405 filed Dec. 28, 1984, incorporated herein by reference.

The zeolite catalysts preferred for use herein include the crystalline aluminosilicate zeolites having a silica to alumina ratio of at least 12, a constraint index of about 1 to 12 and acid cracking activity of about 1-200. 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 disclosed and claims in U.S. Pat. No. 3,702,886 and U.S. Pat. No. Re. 29,948; ZSM-11 is disclosed and claimed in U.S. Pat. 3,709,979. Also, see U.S. Pat. No. 3,832,449 for ZSM-12; U.S. Pat. No. 4,076,979. Also, see U.S. Pat. No. 3,832,449 for ZSM-12; U.S. Pat. No. 4,076,842 for ZSM-23; U.S. Pat. No. 4,016,245 for ZSM-35 and U.S. Pat. No. 4,046,839 for ZSM-38. The disclosures of these patents are incorporated herein by reference. A suitable catalyst for oxygenate conversion is HZSM-5 zeolite with alumina binder. These medium pore shape selective catalysts are sometimes known as porontectosilicates or "pentasil" catalysts.

Other catalysts and processes suitable for converting methanol/DME to lower olefins are disclosed in U.S. Pat. No. 4,493,263 (Vogt et al.) and European Patent Application 0081683 (Marosi et al.), and ZSM-45. In addition to the preferred aluminosilicates, the borosilicate, ferrosilicate and "silicalite" materials may be employed. ZSM-5 type catalysts are particularly advantageous because the same material may be employed for dehydration of methanol to DME, conversion to lower olefins and ethylene conversion.

In this description, metric units and parts by weight are employed unless otherwise stated. Various reactor configurations may be used, including fluidized bed catalytic reactors, moving bed and fixed bed reactors.

Referring to FIG. 1, the process feedstock (methanol or DME, for instance) is fed via conduit 1 to the primary MTO stage 10 where it is converted to lower olefins and C.sub.5 + gasoline hydrocarbon plus water by dehydration of the oxygenated feedstock. The primary effluent is cooled by exchanger 12 and byproduct water is recovered by phase separation unit 14 from the cooled effluent. Liquid hydrocarbons L consisting essentially of C.sub.5.sup.+ gasoline range materials are recovered from unit 14 and pumped to the absorber unit 18 at higher (secondary stage) pressure. Optionally, the heavy liquid condensed from MTO liquid may be sent to a fractionation system, as indicated by dashed line, to recover aromatics rich C.sub.6.sup.+ components. Alternatively, MTO heavy liquid can be prefractionated to recover C.sub.9.sup.+ aromatics as a separate stream.

Vapor phase effluent V from the primary stage may be compressed to alkylation reaction pressure in unit 16. Propylene, butylenes and amylenes may be separated from the primary stage effluent by sorption fractionation unit 18 to recover a recycle gas stream containing ethene from the primary stage and an olefinic sorbate stream rich in C.sub.3 + olefins. Purified ethene (ie--90+%) can be obtained from cryogenic separation unit 21. A C.sub.3 -C.sub.4 rich olefinic stream may be further prepared for reaction with isobutane or the like at high pressure and low temperature in contact with liquid phase acidic alkylation catalyst. The C.sub.3 +0 sorbate stream from sorber unit 18 is fractionated in tower 24 to recover C.sub.6 + aromatic and aliphatic components. A portion of the gasoline tower bottom is recycled to the sorber 18 as lean sorbent via line 26. Thus the lower olefins and other C.sub.2 -C.sub.5 aliphatic component are fed via conduit 28 for upgrading by reaction in the alkylation stage 30. Fresh isoparaffin (e.g.-i-C.sub.4) and recycle are fed via conduit 29. Secondary stage alkylation effluent is then separated into C.sub.3.sup.- light gases, isobutane recycle, n-C.sub.4 aliphatics and C.sub.5.sup.+ gasoline and/or light distillate range alkylate hydrocarbons. Advantageously, isobutane is separated from the second stage effluent for recycle to provide a stoichiometric excess and added with fresh feed (i-C.sub.4) to the unit. Optionally, as indicated by dashed line 34, a portion of the liquid paraffin-rich hydrocarbon may be recycled to the interstage sorption unit 18 to supplement the MTO liquids as lean sorbent. The preferred sorption unit is a countercurrent packed tower. Advantageously, the sorbent liquid comprises at least 50-75 wt% of C.sub.6 + hydrocarbons which are the primary stage MTO reaction product. This lean sorbent has excellent properties for selective sorption of the propene, butylene and C.sub.3 + paraffinic components of the primary stage light hydrocarbons. Unfractionated liquid alkylation effluent may be recycled in part, as depicted by dashed line.

In the embodiment of FIG. 2, oxygenate feedstock is converted in MTO unit 110, cooled and passed to separator 114 to provide a C.sub.5 + liquid stream. The light hydrocarbon vapor stream V separated from the primary stage effluent is compressed in a plurality of compression stages to condense liquid olefinic hydrocarbons. The liquid stream is combined with a condensed liquid from succeeding separators. The primary vapor stream is adiabatically compressed by compressor 116A, cooled by exchanger 117 and passed to a succeeding separator 114A, at which point the preceeding phase separation technique is repeated. Likewise other compressor 116B, C separators 114B and 114C operate to provide an ethene-rich recycle stream which is passed via line 113 to sorption fractionation unit 118 and optional cryogenic unit 120. Advantageously, the MTO effluent is received at about atmospheric pressure (eg, 100-150 kPa) and compressed in plural stages to a pressure of about 1500-3000 kPa and separated in the final vessel 114C at about ambient temperature (20.degree.-80.degree. C.). Olefinic liquids rich in C.sub.3.sup.+ aliphatics may be recovered from the final compressor stage and passed with C.sub.5 + liquid hydrocarbon stream to fractionation tower 124 where C.sub.5 + gasoline recycle and product are recovered.

A suitable selective sorption unit is disclosed in U.S. Pat. No. 4,497,968 (Hsia et al), incorporated herein by reference. Ethene-rich vapor withdrawn from the separator 114C can be processed to increase ethene purity to at least 50% in sorption unit 118. This can be achieved by selectively absorbing C.sub.3.sup.+ components in a C.sub.5.sup.+ liquid hydrocarbon sorbent stream comprising aliphatic and aromatic hydrocarbons generated in the MTO unit.

The data in Table I represents a material balance and absorber operating conditions for a countercurrent contact tower unit design according to FIG. 2 and Runs 1-3. The vertical tower has 19 theoretical stages, with C.sub.6 + gasoline lean oil from the interstage sorbate fractionator being introduced at the top (stage 1), olefin vapor and liquid feed being fed at stages 9 and 10 respectively. Heat of sorption is removed by pumparound cooling at stages 5 and 8. The three runs correspond to different lean oil rates. In the overhead stream, molar units are gm-moles per metric tonne (MT) of methanol charged to the process. The operating conditions are chosen to provide a maximum ethene content of 0.2 mol% in the sorbate.

Runs 4 to 6 (Table II) differ from runs 1 to 3 in that two lean oil compositions are employed. To the top stage is fed gasoline sorbent from the sorbate fractionator and to a lower stage (5) is fed heavy liquid separated from the MTO effluent, according to FIG. 1. Run 7 (Table III) is similar to run 4 except for sorbent rate and 1.5% ethene in the sorbate. Run 8 is similar to run 1 except for 0.1% ethene in the sorbate. Run 9 is similar to run 7 except for the use of compression section recontact between condensate and vapor prior to the sorption unit.

The alkylation process employed herein for upgrading olefins is a well known industrial technique for reacting alkenes with tertiary alkanes (isoparaffins), such as isobutane, isopentane, isohexane, etc. The resulting product is a C.sub.7.sup.+ branched chain paraffinic material useful as aviation gasoline, jet fuel or the like. The alkylation of paraffins can be carried out either thermally or catalytically; however, acid catalyst is preferred. Thermal or noncatalytic alkylation of a paraffin with an olefin is carried out at high temperatures (about 500.degree. C.) and pressures 21-41 MPa (3000-6000 psi). Under these conditions, both normal and isoparaffins can be brought into reaction by a free-radical mechanism. Thermal alkylation is not known to be practiced commercially.

The catalytic alkylation of paraffins involves the addition of an isoparaffin containing a tertiary hydrogen to an olefin. The process is used in the petroleum industry to prepare highly branched paraffins mainly in the C.sub.7 to C.sub.9 range, that are high-quality fuels. The overall process is complex, requiring control of operating conditions and of catalyst. The process conditions and the product composition depend on the particular hydrocarbons involved.

The preferred processes are those brought about by the conventional protonic and Lewis catalysts. Propene can be brought into reaction with an isoparaffin in the presence of either concentrated sulfuric acid or hydrogen fluoride. The heptanes produced by alkylation of isobutane with propene are mainly 2,3- and 2,4-dimethylpentane. Propene is alkylated preferrably as a component of a C.sub.3 -C.sub.4 fraction. HF catalysts for alkylation of isobutane with 1- and 2-butenes give both dimethylhexanes and trimethylpentanes. The product obtained from alkylation of isobutane with isobutylene at low temperature (e.g., -25.degree. C.) with hydrogen fluoride is 2,2,4-trimethylpentane.

During use the acid catalysts may become diluted with byproduct hydrocarbons and as a result decrease in activity. Sulfuric acid concentrations are maintained at about 90%. Hydrogen fluoride concentrations of 80-90% are common, although the optimum concentration depends on the reaction temperature and reactor geometry. Operation below these acid concentrations generally causes incomplete conversion or polymerization. With sulfuric acid, the product quality is improved when temperatures are reduced to the range of 0.degree.-10.degree. C. Cooling requirements are obtained by low temperature flashing of unreacted isobutane. With hydrogen fluoride, the reaction process is less sensitive to temperature, and temperatures of 0.degree.-40.degree. C. can be used. Some form of heat removal is essential because the heat of reaction is approximately 14.times.10.sup.5 J/kg (600 Btu/lb) of butenes converted. Typically the elevated pressure for alkylation by these acid catalysts is about 1500 to 3000 kPa (200-300 psig).

In order to prevent polymerization of the olefin as charged, an excess of isobutane is present in the reaction zone. Isobutane-to-olefin molar ratios of 6:1 to 14:1 are common, more effective suppression of side reactions being produced by the higher ratios.

The typical alkylation reaction employs a two-phase system with a low solubility of the isobutane in the catalyst phase. In order to ensure intimate contact of reactants and catalyst, efficient mixing is provided. This is important with sulfuric acid because of the low solubility of isobutane in the catalyst phase. In addition, the higher viscosity of the sulfuric acid requires a greater mixing energy to assure good contact. The solubility of the hydrocarbon reactants in the catalyst phase is increased by the presence of the unsaturated organic diluent held by the acid catalyst. This organic diluent also has been considered a source of carbonium ions that promote the alkylation reaction.

For the hydrofluoric acid system, reactive i-C.sub.4 H.sub.8 readily alkylates to give an excellent product. The alkylation of pure 1-C.sub.4 H.sub.8 by itself proceeds with considerable isomerization of the 1-C.sub.4 H.sub.8 to 2-C.sub.4 H.sub.8 followed by alkylation to give a highly branched product. The presence of i-C.sub.4 H.sub.8 accelerates the alkylation reaction and allows less time for olefin isomerization. Consequently the reaction produces an alkylate with a lowered antiknock value. For the sulfuric acid system, i-C.sub.4 H.sub.8 tends to oligomerize and causes other side reaction products of inferior quality; but the isomerization of 1-C.sub.4 H.sub.8 to 2-C.sub.4 H.sub.8 proceeds more completely, thereby favoring formation of superior products. Thus for mixed olefin feeds such as described above, the two factors with both catalyst systems counteract each other to provide products of similar antiknock properties.

The olefin-producing MTO process may simultaneously generate isobutane, but the amount may be insufficient to alkylate the coproduced olefins. A suitable outside source of isobutane is natural gas or a byproduct of methanol-to-gasoline (MTG) processes.

Suitable alkylation processes are described in U.S. Pat. Nos. 3,879,489 (Yurchak et al), 4,115,471 (Kesler), 4,377,721 (Chester) and in the Kirk-Othmer Encyclopedia of Chemical Technology, Vol. 2, pp. 50-58 (3rd Ed., 1978) John Wiley & Sons, incorporated herein by reference.

The combined processes are an effective means for converting oxygenated organic compounds, such as methanol, DME, lower aliphatic ketones, aldehydes, esters, etc, to valuable hydrocarbon products. Thermal integration is achieved by employing heat exchangers between various process streams, towers, absorbers, etc.

Various modifications can be made to the system, especially in the choice of equipment and non-critical processing steps. While the invention has been described by specific examples, there is no intent to limit the inventive concept as set forth in the following claims.