CA1291770C - Conversion of crude methanol to gasoline with extraction - Google Patents
Conversion of crude methanol to gasoline with extractionInfo
- Publication number
- CA1291770C CA1291770C CA000557482A CA557482A CA1291770C CA 1291770 C CA1291770 C CA 1291770C CA 000557482 A CA000557482 A CA 000557482A CA 557482 A CA557482 A CA 557482A CA 1291770 C CA1291770 C CA 1291770C
- Authority
- CA
- Canada
- Prior art keywords
- methanol
- hydrocarbons
- liquid
- feedstock
- extraction
- Prior art date
- Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
- Expired - Lifetime
Links
Classifications
-
- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
- C07C1/20—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only oxygen atoms as heteroatoms
-
- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10L—FUELS NOT OTHERWISE PROVIDED FOR; NATURAL GAS; SYNTHETIC NATURAL GAS OBTAINED BY PROCESSES NOT COVERED BY SUBCLASSES C10G, C10K; LIQUEFIED PETROLEUM GAS; ADDING MATERIALS TO FUELS OR FIRES TO REDUCE SMOKE OR UNDESIRABLE DEPOSITS OR TO FACILITATE SOOT REMOVAL; FIRELIGHTERS
- C10L1/00—Liquid carbonaceous fuels
- C10L1/02—Liquid carbonaceous fuels essentially based on components consisting of carbon, hydrogen, and oxygen only
- C10L1/023—Liquid carbonaceous fuels essentially based on components consisting of carbon, hydrogen, and oxygen only for spark ignition
-
- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C2529/00—Catalysts comprising molecular sieves
- C07C2529/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites, pillared clays
- C07C2529/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
- C07C2529/40—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the pentasil type, e.g. types ZSM-5, ZSM-8 or ZSM-11
-
- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
- Y02P30/00—Technologies relating to oil refining and petrochemical industry
- Y02P30/20—Technologies relating to oil refining and petrochemical industry using bio-feedstock
-
- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y10—TECHNICAL SUBJECTS COVERED BY FORMER USPC
- Y10S—TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y10S203/00—Distillation: processes, separatory
- Y10S203/24—Mobile, marine distiller
Abstract
ABSTRACT
A process for converting crude aqueous methanol feedstock to liquid hydrocarbons in contact with a medium pore shape selective crystalline acid zeolite catalyst, such as HZSM-5. In a preferred embodiment, the novel technique comprises the steps of: (a) contacting a crude methanol feedstock containing a minor amount of water with a liquid hydrocarbon extraction stream 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 extracted methanol substantially free of water to the reaction zone under process conditions to convert substantially all methanol to hydrocarbons; (c) cooling reaction effluent to recover aqueous liquid byproduct stream, gas rich in C2- hydrocarbons, liquid rich in C3-C4 and C5+ hydrocarbons; and (d) contacting crude methanol feedstock with at least a portion of the liquid hydrocarbons employed as extraction liquid.
A process for converting crude aqueous methanol feedstock to liquid hydrocarbons in contact with a medium pore shape selective crystalline acid zeolite catalyst, such as HZSM-5. In a preferred embodiment, the novel technique comprises the steps of: (a) contacting a crude methanol feedstock containing a minor amount of water with a liquid hydrocarbon extraction stream 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 extracted methanol substantially free of water to the reaction zone under process conditions to convert substantially all methanol to hydrocarbons; (c) cooling reaction effluent to recover aqueous liquid byproduct stream, gas rich in C2- hydrocarbons, liquid rich in C3-C4 and C5+ hydrocarbons; and (d) contacting crude methanol feedstock with at least a portion of the liquid hydrocarbons employed as extraction liquid.
Description
`` ~2~770 CON~ERSION OF CRUDE METHANOL TO GASOLIN~
WITH EXTRACTION
This invention relates to a process for converting methanol and/or related oxygenated organic compounds to gasoline boiling range hydrocarbons. In particular, the invention relates to a process for removing water from crude aqueous methanol feedstock or the like prior to catalytic reaction in a MTG (methanol to gasoline) reactor.
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 liqùid 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 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%.
The conversion of methanol and other lower aliphatic oxygenates to hydrocarbon products may take place in a fixed bed process as described in U.S. Patent Nos. 3,998,899; 3,931,349 and 4,035,43û. Fluidized bed catalysis has been described in US Patents 4,251,484 (Daviduk et al) and 4,513,160 (Avidan and Kam). In the fixed bed process, the methanol is usually 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 which are mainly in the range of light gas to 12~7~
F-4250 --~--gasoline. 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 become deactivated by the presence of excess water vapor at the reaction temperatures employed; but this step is not essential.
In the operation of an adiabatic fixed bed process, a major problem is thermal balance. The conversion of the oxygenated feed stream (methanol, DME) to the hydrocarbons is a strongly exothermic reaction liberating approximately 1480 kl. (1400 Btu) of heat per kilogram of methanol. In an uncontrolled adiabatic reactor this would result in a temperature rise which would lead to extremely fast catalyst aging rates or even to damage to the catalyst.
Furthermore, the high temperatures which might occur could cause undesirable products to be produced or the product distribution could be unfavorably changed. It is therefore necessary that some method should be provided to maintain the catalyst bed within desired temperature limits by dissipating the heat of the reaction.
One method is to employ a light gas portion of the hydrocarbon product as recycle, as described in U.S. Patent No.
3,931,349 (Kuo). Typically, cooled light hydrocarbon gas, rich in methane, ethane, etc., is separated from the C5+ gasoline and C3-C4 LPG products, re-compressed and reheated before being mixed with the reactant feedstream entering the bed of conversion catalyst. Although effective in controlling bed temperature, the expense of cooling the recycle gas, compressing `it and re-heating it add to the cost of the convers`ion.
Typically the crude methanol feestock employed in MTG
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 12~
distillation system described by Mao et al in U.S. Patent 4,665,249.
According to the present invention, an improved process has been found for converting crude aqueous methanol feedstock to gasoline boiling range hydrocarbons in contact with a medium pore shape selective crystalline acid zeolite catalyst. The improved process comprises contacting the aqueous methanol feedstock with a liquid hydrocarbon extractant under liquid extraction conditions, recovering an aqueous phase containing the major amount of water and a portion of methanol introduced with the feedstock, recovering an organic extract phase comprising the hydrocarbon extractant and the methanol introduced in the aqueous feedstock; and converting the extracted methanol at elevated temperature under catalytic reaction conditions tp produce predominantly gasoline range liquid hydrocarbons and a minor amount of C3-C4 hydrocarbons. The li~uid hydrocarbons and/or C3-C4 may be passed to the extraction step to provide a selective composition for removing methanol from the crude feedstream. This technique is adapted for continuous process conditions.
In the drawings, FIG. 1 is a schematic process flowsheet depicting the present invention;
FIG.2 is a flow diaaram of an extraction and methanol dewatering subsystem using C3 alkane extractant;
FIG.3 is a graphic plot of methanol recovery vs. carbon ; 25 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 schematic process flowsheet for an alternative embodi~ent of the invention using full ranae C3~ liquid extractant.
The present Process is useful for the conversion of methanol. The hy~rocarbon product will he a hydrocarbon mixture ~291~70 ranging from light gas to heavier fractions, but will nenerally be concentrated in the gasoline boiling range (C5-220C). The MTG
process is particularly useful in the catalytic conversion of methanol and its corresponding ether to normally liquid hydrocarbons in the C5-C10 range. For convenience, the process will be described below with reference to such a specific methanol/DME
conversion process; although it should be remembered that the principles are applicable to a broader range of conversions.
When methanol is used as the starting material for the process it is preferred to subject it to an initial dehydration step to form an intermediate product comprising dimethyl ether (DME).
The DME is then passed to the hydrocarbon step with either complete, partial or no separation of the unreacted methanol and the water produced as a by-product of the dehydration. However, it is not essential to carry out this dehydration even though it is preferred.
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 is withdrawn via conduit 16. The lighter organic extract phase containing hydrocarbon extraction solvent and a significant amount of feedstock methanol is recovered from extraction unit 12 via conduit 18, pressurized by pump means 20, passed through heat exchanger 22, furnace 24 and introduced at elevated temperature and process conditions suitable for conversion of substantially all of the methanol in contact with the MTG catalyst in reactor system 30.
The crude feedstock mixture of methanol may be fed directly to a fixed bed MTG (methanol to gasoline) catalytic reactor. From the reactor system 30, the effluent product stream leaves via line 32 to a heat exchanger condenser 34, and is passed as a three-phase mixture to a primary phase separator 36. The temperature and pressure of the separator are maintained such that the products are split into three separate streams. The byproduct water is recovered via line 38. The light hydrocarbon aas phase is passed via line 39 to a deethanized tower or similar gas separation unit for recovery f C2 fuel gas byproduct and LPG rich in C3-C4 hydrocarbons, especially propane and butanes, which may be further fractionated to recover isobutane or other components thereof. The condensed liquid hydrocarbon stream from separator unit 36 is further fractionated in tower 44 to provide the C5-C8 gasoline product stream, a C9+ aromatics rich heavy gasoline stream and a C4- overhead vapor stream, which is combined with vapor stream 39 flowing to the deethanizer tower 40. A recycle C3-C4 stream including at lEast a portion of LPG hydrocarbons recovered from the process reactor effluent is sent via conduit 14 to use as extractant liquid , along with any added propane or other makeup lS 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 in tower 50 to provide a methanol-rich auxiliary feedstream 50A. Alternatively this may be combined with crude feedstock in conduit 10.
EXTRACTION UNIT OPERATION
Extraction conditions may be varied depending upon water content of crude methanol feedstock and the required water-oxygenate separation. A typical crude feedstock material is methanol containing 4 to 17% by weight water. The extraction unit 12 can be designed as a stirred multi-stage vertical extraction column adapted for continuous operation at elevated pressure. Any suitable extraction equipment may be employed, including cocurrent, cross-current or sinale stage contactors, wherein the liquid methanol feedstock is intimately contacted with a substantially immiscible liquid hydrocarbon solvent, which may be a mixture of C3 aliphatic components or relatively pure propane, butane, etc. This unit operation is described in Kirk-Othmer Encyclopedia 1291~71~
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 s 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.
WITH EXTRACTION
This invention relates to a process for converting methanol and/or related oxygenated organic compounds to gasoline boiling range hydrocarbons. In particular, the invention relates to a process for removing water from crude aqueous methanol feedstock or the like prior to catalytic reaction in a MTG (methanol to gasoline) reactor.
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 liqùid 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 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%.
The conversion of methanol and other lower aliphatic oxygenates to hydrocarbon products may take place in a fixed bed process as described in U.S. Patent Nos. 3,998,899; 3,931,349 and 4,035,43û. Fluidized bed catalysis has been described in US Patents 4,251,484 (Daviduk et al) and 4,513,160 (Avidan and Kam). In the fixed bed process, the methanol is usually 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 which are mainly in the range of light gas to 12~7~
F-4250 --~--gasoline. 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 become deactivated by the presence of excess water vapor at the reaction temperatures employed; but this step is not essential.
In the operation of an adiabatic fixed bed process, a major problem is thermal balance. The conversion of the oxygenated feed stream (methanol, DME) to the hydrocarbons is a strongly exothermic reaction liberating approximately 1480 kl. (1400 Btu) of heat per kilogram of methanol. In an uncontrolled adiabatic reactor this would result in a temperature rise which would lead to extremely fast catalyst aging rates or even to damage to the catalyst.
Furthermore, the high temperatures which might occur could cause undesirable products to be produced or the product distribution could be unfavorably changed. It is therefore necessary that some method should be provided to maintain the catalyst bed within desired temperature limits by dissipating the heat of the reaction.
One method is to employ a light gas portion of the hydrocarbon product as recycle, as described in U.S. Patent No.
3,931,349 (Kuo). Typically, cooled light hydrocarbon gas, rich in methane, ethane, etc., is separated from the C5+ gasoline and C3-C4 LPG products, re-compressed and reheated before being mixed with the reactant feedstream entering the bed of conversion catalyst. Although effective in controlling bed temperature, the expense of cooling the recycle gas, compressing `it and re-heating it add to the cost of the convers`ion.
Typically the crude methanol feestock employed in MTG
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 12~
distillation system described by Mao et al in U.S. Patent 4,665,249.
According to the present invention, an improved process has been found for converting crude aqueous methanol feedstock to gasoline boiling range hydrocarbons in contact with a medium pore shape selective crystalline acid zeolite catalyst. The improved process comprises contacting the aqueous methanol feedstock with a liquid hydrocarbon extractant under liquid extraction conditions, recovering an aqueous phase containing the major amount of water and a portion of methanol introduced with the feedstock, recovering an organic extract phase comprising the hydrocarbon extractant and the methanol introduced in the aqueous feedstock; and converting the extracted methanol at elevated temperature under catalytic reaction conditions tp produce predominantly gasoline range liquid hydrocarbons and a minor amount of C3-C4 hydrocarbons. The li~uid hydrocarbons and/or C3-C4 may be passed to the extraction step to provide a selective composition for removing methanol from the crude feedstream. This technique is adapted for continuous process conditions.
In the drawings, FIG. 1 is a schematic process flowsheet depicting the present invention;
FIG.2 is a flow diaaram of an extraction and methanol dewatering subsystem using C3 alkane extractant;
FIG.3 is a graphic plot of methanol recovery vs. carbon ; 25 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 schematic process flowsheet for an alternative embodi~ent of the invention using full ranae C3~ liquid extractant.
The present Process is useful for the conversion of methanol. The hy~rocarbon product will he a hydrocarbon mixture ~291~70 ranging from light gas to heavier fractions, but will nenerally be concentrated in the gasoline boiling range (C5-220C). The MTG
process is particularly useful in the catalytic conversion of methanol and its corresponding ether to normally liquid hydrocarbons in the C5-C10 range. For convenience, the process will be described below with reference to such a specific methanol/DME
conversion process; although it should be remembered that the principles are applicable to a broader range of conversions.
When methanol is used as the starting material for the process it is preferred to subject it to an initial dehydration step to form an intermediate product comprising dimethyl ether (DME).
The DME is then passed to the hydrocarbon step with either complete, partial or no separation of the unreacted methanol and the water produced as a by-product of the dehydration. However, it is not essential to carry out this dehydration even though it is preferred.
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 is withdrawn via conduit 16. The lighter organic extract phase containing hydrocarbon extraction solvent and a significant amount of feedstock methanol is recovered from extraction unit 12 via conduit 18, pressurized by pump means 20, passed through heat exchanger 22, furnace 24 and introduced at elevated temperature and process conditions suitable for conversion of substantially all of the methanol in contact with the MTG catalyst in reactor system 30.
The crude feedstock mixture of methanol may be fed directly to a fixed bed MTG (methanol to gasoline) catalytic reactor. From the reactor system 30, the effluent product stream leaves via line 32 to a heat exchanger condenser 34, and is passed as a three-phase mixture to a primary phase separator 36. The temperature and pressure of the separator are maintained such that the products are split into three separate streams. The byproduct water is recovered via line 38. The light hydrocarbon aas phase is passed via line 39 to a deethanized tower or similar gas separation unit for recovery f C2 fuel gas byproduct and LPG rich in C3-C4 hydrocarbons, especially propane and butanes, which may be further fractionated to recover isobutane or other components thereof. The condensed liquid hydrocarbon stream from separator unit 36 is further fractionated in tower 44 to provide the C5-C8 gasoline product stream, a C9+ aromatics rich heavy gasoline stream and a C4- overhead vapor stream, which is combined with vapor stream 39 flowing to the deethanizer tower 40. A recycle C3-C4 stream including at lEast a portion of LPG hydrocarbons recovered from the process reactor effluent is sent via conduit 14 to use as extractant liquid , along with any added propane or other makeup lS 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 in tower 50 to provide a methanol-rich auxiliary feedstream 50A. Alternatively this may be combined with crude feedstock in conduit 10.
EXTRACTION UNIT OPERATION
Extraction conditions may be varied depending upon water content of crude methanol feedstock and the required water-oxygenate separation. A typical crude feedstock material is methanol containing 4 to 17% by weight water. The extraction unit 12 can be designed as a stirred multi-stage vertical extraction column adapted for continuous operation at elevated pressure. Any suitable extraction equipment may be employed, including cocurrent, cross-current or sinale stage contactors, wherein the liquid methanol feedstock is intimately contacted with a substantially immiscible liquid hydrocarbon solvent, which may be a mixture of C3 aliphatic components or relatively pure propane, butane, etc. This unit operation is described in Kirk-Othmer Encyclopedia 1291~71~
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 s 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~ is fed via inlet 110 to extraction unit 112 at a flow rate of 16364 moles MeOH + 6080 moles H20, where it is contacted with 89776 moles of C3 (propane) hydrocarbon at a temperature of 43C. The extract phase containing 9596 moles of MeOH, 325 moles H20 (less than 2 wt~), and 89560 moles C3 hydrocarbon is passed via conduit 118, heat exchanger 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 H20, 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.
The solubility of methanol in hydrocarbons coupled with the insolubility of hydrocarbons in water makes extraction a viable step in methanol dewatering. The heavy liquid raffinate phase, consisting mostly of water and methanol (less than 4 wt. ~
hydrocarbon), is distilled to recover unextracted MeOH and partitioned hydrocarbon By implementing extraction first, the feed to the distillation column is significantly reduced, thus reducing tower diameter, reboiler duty and overhead condenser duty.
For this example propane is used as the hydrocarbon, as is present in a 4:1 molar ratio. Liquid propane is contacted in a liqui~/liquid phase separator with crude methanol at about 43C
(110F). The estimated values of methanol recovered and water entrained are determined according to FIGS 3 and 4. This improved technique is compared to a conventional distillation of crude methanol.
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Raffinate stream 116, containing 6768 moles MeOH, 5755 moles H20, 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.
The solubility of methanol in hydrocarbons coupled with the insolubility of hydrocarbons in water makes extraction a viable step in methanol dewatering. The heavy liquid raffinate phase, consisting mostly of water and methanol (less than 4 wt. ~
hydrocarbon), is distilled to recover unextracted MeOH and partitioned hydrocarbon By implementing extraction first, the feed to the distillation column is significantly reduced, thus reducing tower diameter, reboiler duty and overhead condenser duty.
For this example propane is used as the hydrocarbon, as is present in a 4:1 molar ratio. Liquid propane is contacted in a liqui~/liquid phase separator with crude methanol at about 43C
(110F). The estimated values of methanol recovered and water entrained are determined according to FIGS 3 and 4. This improved technique is compared to a conventional distillation of crude methanol.
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In both cases, the same crude methanol stream is dewatered. In both cases the same amount of methanol and water is recovered (4 wt. %
water on a hydrocarbon-free basis), and the feed temperature to the distillation tower was kept constant.
Table 1 shows estimated methanol recovery and water entrainment for 5 hydrocarbon solvents, at different extraction temperatures for various hydrocarbon:crude methanol ratios. These data shown graphically in FIGS. 3-4, indicate better methanol recovery with lower molecular weight hydrocarbons. At higher temperatures, recovery increases, but so does water entrainment.
When two liquid phases are present, pressure appears to have no significant effect.
Table 2 indicates that reboiler duty decreases by 51%, overhead condenser duty decreases by 56~, and the distillation tower diameter decreases by 41% when extraction is used prior to distillation. Preheat of the feed or recycle C3 hydrocarbon can occur even if the extraction is not performed, so it is not a dewatering process energy expense.
` lZ91770 METHANOL DISTILLATION TOWER OPERATING PARAMETERS
Distillation Extraction/
Only Distillation Feed Methanol (lbmol/hr) 16364. 6768.
Water 6080. 5755.
Propane 216.
Overhead Product Methanol(mol/hr)16357.10 6761 Water 1181.22 850 Propane 216 Bottoms Product Methanol(mbl/hr) 6.90 6.93 Water 4898.78 4905 Duties Condenser, 446 (423) 196 (186) million kJ/hr (million BTu/hr) Reboiler 381 (361) 187 (177) : : REACTOR SYSTEM PROCESS CONDITIONS
The feedstock for a fixed bed MTG catalytic reactor can be the effluent from a DME (dimethyl ether) dehydration reactor. In such a case, the intermediate stream is an equilibrium mixture of : 5 methanol, water and DME. If desired, complete or partial separation into constituent product streams of water, methanol, and DME may be carried out as described in U.S. Patent No. 4,035,430 by : condensation or fractionation, depending upon the degree of purity i291770 desired~ Removal of at least the water is desirable because the conversion catalysts used in the fixed bed catalytic reactor usually become deactivated under hydrothermal reaction conditions encountered in the conversion. Although it is not feasible to eliminate the presence of water vapor completely from the conversion, because water is a by-product; the removal of water from the charge will lead to a reduction of the water vapor partial pressure in the conversion reactor and will, accordingly, lead to an increase in the useful life of the catalyst.
The dehydration products are passed to the fixed bed MTG
reactor under conversion conditions at elevated temperature and pressure. The conversion to hydrocarbons is preferably catalyzed by a medium pore crystalline metallosilicate zeolite catalyst having Bronsted acid functionality. Recent developments in zeolite lS 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 A1, 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 cystalline 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. 3,702,866 (Argauer, et al.).
The preferred class of catalysts is characterized by a Constraint Index of 1 to 12 and a silica:alumina ratio of at least 12:1 and preferably higher e.g., 20:1 to 100:1. Preferred aluminosilicate zeolites which have the specified values of Constraint Index and silica:alumina ratio include zeolites ZSM-5, ZSM-ll, ZSM-12, ZSM-35, ZSM-38 and ZSM-48, which are described in U.S. Patent Nos. 3,702,886 (ZSM~5), 3,709,979 (ZSM-ll), 3,832,449 (ZSM-12), 4,076,842 (ZSM-23) and 4,016,245 (ZSM-35),and 4,046,859 (ZSM-38). The zeolite catalyst used is at i2917f'70 least partly in the hydrogen form e.g., HZSM-5; but other cations, e.g., Periodic Groups III-VIII or rare earth cations, may also be present. Other cations e.g., metal cations can be introduced by conventional base exchange technioues.
In this invention, the exothermic reaction can be moderated by recycling C3-C4 hydrocarbons employed in extracting the feedstock methanol. Rather than employing a C2 recycled light hydrocarbon gas as the diluent, C3-C4 rich (LPG) recycle fraction is used. The LPG stream may comprise, on a molar basis, about 48% isobutane, about 16% propane, and about 18~ n-butane, with a minor amount of olefins. The MTG reactor can be operated under process conditions that convert propane or other alkanes. Where a supply of outside propane is available, a major amount of the extractant can be comprised of recycled and makeup propane. Typically, about 95% of the LPG recycle are C3-C7 saturated hydrocarbons.
A major advantage of employing an LPG recycle stream is that liquid can be repumped economically in order to recycle it back to the fixed bed MTG catalytic reactor. If light hydrocarbon gas is used for the recycle, it must be recompressed at great expense. A
large gas recycle compressor usually requires a source of high pressure superheated steam; whereas the C3-C4 recycle design demands much less energy to achieve the process reaction pressure.
Another advantage of an LPG recycle versus a light gas recycle is that an increased circulation of LPG range material through the reaction zone can provide increased conversion of C3-C4 olefins to gasoline range hydrocarbons. The overall MTG gasoline yield would then be greater. In addition, improved selectivity toward reduced fuel gas production is obtained.
In the embodiment of Fig. 1, extracted methanol is passed to the reactor system with hydrocarbon extractant containinn LPG
recycle. The combined feedstock and recycle is conducted at a temperature of about 340-345C (650~F) and a pressure of about 1000-1100 kPa (154 psia) to a fixed bed MTG catalytic reactor.
Effluent from the fixed bed MTG reaction zone is passed to a heat exchange condenser and then to a separation zone without significant loss of pressure. In the separation zone the effluent stream is partitioned into three streams. Typically, the separator is operated at conditions of about 35-40C (100F) and about 700-750 kPa (107 psia). From the separator the liquid hydrocarbon fraction is passed to a pump where it is repressurized to about 1100 kPa to 1400 kPa. From the heat exchange evaporator the partially vaporized fraction is passed to an accumulator. The accumulator is operated at about 90-95C (195F) and about 1200-1300 kPa (186 psia).
From the accumulator a vaporized fraction containing a major amount of C3-C4 saturated hydrocarbons (LPG) is passed to a zone where it is separated from LPG make. The purified LPG fraction is then conducted to a heat exchange evaporator where it is brought to a temperature of about 335C (636F) before or after mixing with feedstock.
Referring to Fig. 5 of the drawing , crude methanol (MeOH) feedstock is introduced via conduit 210 and pump 212, heated in feedstock exchanger 214, and passed to a first stage DME reactor 216, where the aqueous methanol feed is contacted at elevated temperature with a dehydration catalyst, such as gamma-alumina to provide an equilibrium mixture of methanol, water and dimethylether. After cooling this first effluent stream in heat exchanger 218, the liquid oxygenate-rich stream is contacted with a liquid hydrocarbon extractant stream rich in C3-C10 hydrocarbons (especially C3-C7 alkanes) in extraction unit 22n, thereby transferring a major portion of the methanol and dimethyl ether oxygenates to the organic extract phase. The aaueous raffinate phase is recovered from extractor 220 and passed to distillation unit 222 for concentrating the oxyaenates prior to recycling with feedstream 210. The lighter organic extract phase containing hy~rocarbon extraction solvent and the major amount of feedstock .~.f~ ~
i291770 methanol and dimethyl ether is recovered from extraction unit 220 and pressurized by pump means 224, passed through heat exchanger 226, and phase separator unit 228 to vaporize the oxygenates and a portion of volatile C3-C4 hydrocarbons. The oxygenate-rich vapor stream containing C3-C4 diluent is then heated in exchanger 229 and introduced at elevated temperature and process conditions suitable for conversion of substantially all of the oxygenates in contact with the MTG catalyst in reactor system 230. From the reactor system 230, the effluent product stream is passed to heat exchange condenser 234, and is passed as a three-phase mixture to a phase separator 236. The temperature and pressure of the separator are maintained such that the products are split into three separate streams. The MTG byproduct water is recovered via line 238. The light hydrocarbon gas phase is removed via line 239. The condensed liquid hydrocarbon stream from separator unit 236 is then contacted with oxygenates in extractor unit 220 to form the organic extract stream passed to phase separator 228 and further fractionated in tower 244 to provide the C5+ MTG gasoline product stream.
While the invention has been described by particular examples, there is no intent to limit the inventive concept except as set forth in the following claims:
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â~
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a~ ~, o 0 ~ 0 o ~ ~ t~l 0 U~ ,~ o 1~ ~ ~
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S F ~ ~ O 0 ~ ~O 0 ~ ~O 0 ~\ ~O 0 a~ I_ ~ ~ O 0 ~t ~O 0 ~ O 0 c o a~
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1~ LLI C~ ~n ~ C C C C
~291~
In both cases, the same crude methanol stream is dewatered. In both cases the same amount of methanol and water is recovered (4 wt. %
water on a hydrocarbon-free basis), and the feed temperature to the distillation tower was kept constant.
Table 1 shows estimated methanol recovery and water entrainment for 5 hydrocarbon solvents, at different extraction temperatures for various hydrocarbon:crude methanol ratios. These data shown graphically in FIGS. 3-4, indicate better methanol recovery with lower molecular weight hydrocarbons. At higher temperatures, recovery increases, but so does water entrainment.
When two liquid phases are present, pressure appears to have no significant effect.
Table 2 indicates that reboiler duty decreases by 51%, overhead condenser duty decreases by 56~, and the distillation tower diameter decreases by 41% when extraction is used prior to distillation. Preheat of the feed or recycle C3 hydrocarbon can occur even if the extraction is not performed, so it is not a dewatering process energy expense.
` lZ91770 METHANOL DISTILLATION TOWER OPERATING PARAMETERS
Distillation Extraction/
Only Distillation Feed Methanol (lbmol/hr) 16364. 6768.
Water 6080. 5755.
Propane 216.
Overhead Product Methanol(mol/hr)16357.10 6761 Water 1181.22 850 Propane 216 Bottoms Product Methanol(mbl/hr) 6.90 6.93 Water 4898.78 4905 Duties Condenser, 446 (423) 196 (186) million kJ/hr (million BTu/hr) Reboiler 381 (361) 187 (177) : : REACTOR SYSTEM PROCESS CONDITIONS
The feedstock for a fixed bed MTG catalytic reactor can be the effluent from a DME (dimethyl ether) dehydration reactor. In such a case, the intermediate stream is an equilibrium mixture of : 5 methanol, water and DME. If desired, complete or partial separation into constituent product streams of water, methanol, and DME may be carried out as described in U.S. Patent No. 4,035,430 by : condensation or fractionation, depending upon the degree of purity i291770 desired~ Removal of at least the water is desirable because the conversion catalysts used in the fixed bed catalytic reactor usually become deactivated under hydrothermal reaction conditions encountered in the conversion. Although it is not feasible to eliminate the presence of water vapor completely from the conversion, because water is a by-product; the removal of water from the charge will lead to a reduction of the water vapor partial pressure in the conversion reactor and will, accordingly, lead to an increase in the useful life of the catalyst.
The dehydration products are passed to the fixed bed MTG
reactor under conversion conditions at elevated temperature and pressure. The conversion to hydrocarbons is preferably catalyzed by a medium pore crystalline metallosilicate zeolite catalyst having Bronsted acid functionality. Recent developments in zeolite lS 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 A1, 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 cystalline 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. 3,702,866 (Argauer, et al.).
The preferred class of catalysts is characterized by a Constraint Index of 1 to 12 and a silica:alumina ratio of at least 12:1 and preferably higher e.g., 20:1 to 100:1. Preferred aluminosilicate zeolites which have the specified values of Constraint Index and silica:alumina ratio include zeolites ZSM-5, ZSM-ll, ZSM-12, ZSM-35, ZSM-38 and ZSM-48, which are described in U.S. Patent Nos. 3,702,886 (ZSM~5), 3,709,979 (ZSM-ll), 3,832,449 (ZSM-12), 4,076,842 (ZSM-23) and 4,016,245 (ZSM-35),and 4,046,859 (ZSM-38). The zeolite catalyst used is at i2917f'70 least partly in the hydrogen form e.g., HZSM-5; but other cations, e.g., Periodic Groups III-VIII or rare earth cations, may also be present. Other cations e.g., metal cations can be introduced by conventional base exchange technioues.
In this invention, the exothermic reaction can be moderated by recycling C3-C4 hydrocarbons employed in extracting the feedstock methanol. Rather than employing a C2 recycled light hydrocarbon gas as the diluent, C3-C4 rich (LPG) recycle fraction is used. The LPG stream may comprise, on a molar basis, about 48% isobutane, about 16% propane, and about 18~ n-butane, with a minor amount of olefins. The MTG reactor can be operated under process conditions that convert propane or other alkanes. Where a supply of outside propane is available, a major amount of the extractant can be comprised of recycled and makeup propane. Typically, about 95% of the LPG recycle are C3-C7 saturated hydrocarbons.
A major advantage of employing an LPG recycle stream is that liquid can be repumped economically in order to recycle it back to the fixed bed MTG catalytic reactor. If light hydrocarbon gas is used for the recycle, it must be recompressed at great expense. A
large gas recycle compressor usually requires a source of high pressure superheated steam; whereas the C3-C4 recycle design demands much less energy to achieve the process reaction pressure.
Another advantage of an LPG recycle versus a light gas recycle is that an increased circulation of LPG range material through the reaction zone can provide increased conversion of C3-C4 olefins to gasoline range hydrocarbons. The overall MTG gasoline yield would then be greater. In addition, improved selectivity toward reduced fuel gas production is obtained.
In the embodiment of Fig. 1, extracted methanol is passed to the reactor system with hydrocarbon extractant containinn LPG
recycle. The combined feedstock and recycle is conducted at a temperature of about 340-345C (650~F) and a pressure of about 1000-1100 kPa (154 psia) to a fixed bed MTG catalytic reactor.
Effluent from the fixed bed MTG reaction zone is passed to a heat exchange condenser and then to a separation zone without significant loss of pressure. In the separation zone the effluent stream is partitioned into three streams. Typically, the separator is operated at conditions of about 35-40C (100F) and about 700-750 kPa (107 psia). From the separator the liquid hydrocarbon fraction is passed to a pump where it is repressurized to about 1100 kPa to 1400 kPa. From the heat exchange evaporator the partially vaporized fraction is passed to an accumulator. The accumulator is operated at about 90-95C (195F) and about 1200-1300 kPa (186 psia).
From the accumulator a vaporized fraction containing a major amount of C3-C4 saturated hydrocarbons (LPG) is passed to a zone where it is separated from LPG make. The purified LPG fraction is then conducted to a heat exchange evaporator where it is brought to a temperature of about 335C (636F) before or after mixing with feedstock.
Referring to Fig. 5 of the drawing , crude methanol (MeOH) feedstock is introduced via conduit 210 and pump 212, heated in feedstock exchanger 214, and passed to a first stage DME reactor 216, where the aqueous methanol feed is contacted at elevated temperature with a dehydration catalyst, such as gamma-alumina to provide an equilibrium mixture of methanol, water and dimethylether. After cooling this first effluent stream in heat exchanger 218, the liquid oxygenate-rich stream is contacted with a liquid hydrocarbon extractant stream rich in C3-C10 hydrocarbons (especially C3-C7 alkanes) in extraction unit 22n, thereby transferring a major portion of the methanol and dimethyl ether oxygenates to the organic extract phase. The aaueous raffinate phase is recovered from extractor 220 and passed to distillation unit 222 for concentrating the oxyaenates prior to recycling with feedstream 210. The lighter organic extract phase containing hy~rocarbon extraction solvent and the major amount of feedstock .~.f~ ~
i291770 methanol and dimethyl ether is recovered from extraction unit 220 and pressurized by pump means 224, passed through heat exchanger 226, and phase separator unit 228 to vaporize the oxygenates and a portion of volatile C3-C4 hydrocarbons. The oxygenate-rich vapor stream containing C3-C4 diluent is then heated in exchanger 229 and introduced at elevated temperature and process conditions suitable for conversion of substantially all of the oxygenates in contact with the MTG catalyst in reactor system 230. From the reactor system 230, the effluent product stream is passed to heat exchange condenser 234, and is passed as a three-phase mixture to a phase separator 236. The temperature and pressure of the separator are maintained such that the products are split into three separate streams. The MTG byproduct water is recovered via line 238. The light hydrocarbon gas phase is removed via line 239. The condensed liquid hydrocarbon stream from separator unit 236 is then contacted with oxygenates in extractor unit 220 to form the organic extract stream passed to phase separator 228 and further fractionated in tower 244 to provide the C5+ MTG gasoline product stream.
While the invention has been described by particular examples, there is no intent to limit the inventive concept except as set forth in the following claims:
Claims (8)
1. A continuous process for converting crude methanol to gasoline boiling range hydrocarbons in a catalytic reaction zone with a crystalline acid zeolite catalyst at elevated temperature and pressure comprising the steps of:
(a) contacting a crude methanol feedstock containing a minor amount of water with a liquid hydrocarbon extraction stream 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 extracted methanol substantially free of water to a reaction zone under process conditions to convert substantially all methanol to hydrocarbons;
(c) cooling reaction effluent to recover aqueous liquid byproduct stream, gas rich in C2- hydrocarbons, liquid rich in C3-C4 and C? hydrocarbons; and (d) contacting crude methanol feedstock with at least a portion of the liquid hydrocarbons employed as extraction liquid.
(a) contacting a crude methanol feedstock containing a minor amount of water with a liquid hydrocarbon extraction stream 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 extracted methanol substantially free of water to a reaction zone under process conditions to convert substantially all methanol to hydrocarbons;
(c) cooling reaction effluent to recover aqueous liquid byproduct stream, gas rich in C2- hydrocarbons, liquid rich in C3-C4 and C? hydrocarbons; and (d) contacting crude methanol feedstock with at least a portion of the liquid hydrocarbons employed as extraction liquid.
2. The process of claim 1 wherein the catalyst consists essentially of ZSM-5 type zeolite.
3. The process of claim 1 or 2 wherein the feedstock consists essentially of methanol and about 4 to 20 wt% water, and wherein the extraction liquid comprises a major amount of propane.
4. The process of claim 1 or 2 wherein a major amount of C3-C4 hydrocarbons are recycled with extracted methanol to the reaction zone as a vapor stream; wherein methanol is separated from the aqueous raffinate and passed to the reaction zone for conversion with the extracted methanol.
5. In a process for converting crude aqueous methanol-containing oxygenate feedstock to gasoline boiling range hydrocarbons in contact with a medium pore shape selective crystalline acid zeolite catalyst;
the improvement which comprises:
contacting the aqueous oxygenate feedstock with a liquid 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 major amount of oxygenate introduced in the feedstock; and converting the extracted oxygenate at elevated temperature under catalytic reaction conditions to produce predominantly gasoline range hydrocarbons and a minor amount of C3-C4 hydrocarbons.
the improvement which comprises:
contacting the aqueous oxygenate feedstock with a liquid 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 major amount of oxygenate introduced in the feedstock; and converting the extracted oxygenate at elevated temperature under catalytic reaction conditions to produce predominantly gasoline range hydrocarbons and a minor amount of C3-C4 hydrocarbons.
6. A process for converting crude aqueous methanol feedstock to liquid hydrocarbons in contact with a medium pore shape selective crystalline acid zeolite catalyst which process comprises the steps of:
catalytically dehydrating methanol feedstock to produce an aqueous mixture of methanol and dimethylether oxygenates;
contacting the aqueous oxygenate mixture with a liquid hydrocarbon extractant comprising C3-C4 hydrocarbons under liquid extraction conditions;
recovering an aqueous raffinate phase containing the major amount of water introduced with the feedstock and made during the dehydrating reaction;
recovering an organic extract phase comprising the hydrocarbon extractant and a major amount of methanol and dimethylether introduced in the aqueous mixture;
pressurizing and heating substantially the entire extract phase to provide a reactant stream comprising hot methanol, dimethylether and C3-C4 hydrocarbon vapor;
converting the extracted methanol and dimethylether in the presence of the C3-C4 hydrocarbon vapor at elevated temperature under catalytic reaction conditions to produce predominantly gasoline range hydrocarbons and a minor amount of C3-C4 hydrocarbons;
separating conversion reaction effluent to recover a liquid hydrocarbon stream rich in C3-C4 and gasoline range hydrocarbons, light gas byproduct, water byproduct; and extracting the aqueous oxygenate mixture with at least a portion of the liquid hydrocarbon stream recovered from reaction effluent as extractant.
catalytically dehydrating methanol feedstock to produce an aqueous mixture of methanol and dimethylether oxygenates;
contacting the aqueous oxygenate mixture with a liquid hydrocarbon extractant comprising C3-C4 hydrocarbons under liquid extraction conditions;
recovering an aqueous raffinate phase containing the major amount of water introduced with the feedstock and made during the dehydrating reaction;
recovering an organic extract phase comprising the hydrocarbon extractant and a major amount of methanol and dimethylether introduced in the aqueous mixture;
pressurizing and heating substantially the entire extract phase to provide a reactant stream comprising hot methanol, dimethylether and C3-C4 hydrocarbon vapor;
converting the extracted methanol and dimethylether in the presence of the C3-C4 hydrocarbon vapor at elevated temperature under catalytic reaction conditions to produce predominantly gasoline range hydrocarbons and a minor amount of C3-C4 hydrocarbons;
separating conversion reaction effluent to recover a liquid hydrocarbon stream rich in C3-C4 and gasoline range hydrocarbons, light gas byproduct, water byproduct; and extracting the aqueous oxygenate mixture with at least a portion of the liquid hydrocarbon stream recovered from reaction effluent as extractant.
7. The process of Claim 6 wherein the extraction step is conducted in a continuous extraction unit having means for contacting the extractant and feedstock under continuous countercurrent conditions.
8. The process of Claim 6 or 7 wherein the catalyst consists essentially of HZSM-5 and wherein the crude feedstock contains about 2 to 30 wt % water.
2490h/0520H
2490h/0520H
Applications Claiming Priority (2)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| US043,718 | 1987-04-29 | ||
| US07/043,718 US4709113A (en) | 1987-04-29 | 1987-04-29 | Conversion of crude methanol to gasoline with extraction |
Publications (1)
| Publication Number | Publication Date |
|---|---|
| CA1291770C true CA1291770C (en) | 1991-11-05 |
Family
ID=21928536
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA000557482A Expired - Lifetime CA1291770C (en) | 1987-04-29 | 1988-01-27 | Conversion of crude methanol to gasoline with extraction |
Country Status (7)
| Country | Link |
|---|---|
| US (1) | US4709113A (en) |
| EP (1) | EP0289234A1 (en) |
| JP (1) | JPS63275691A (en) |
| AU (1) | AU609771B2 (en) |
| CA (1) | CA1291770C (en) |
| NZ (1) | NZ223387A (en) |
| ZA (1) | ZA88693B (en) |
Families Citing this family (28)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US5028400A (en) * | 1987-04-29 | 1991-07-02 | Mobil Oil Corporation | Feedstock preparation and reactor system for conversion to oxygenates to olefins |
| US4777321A (en) * | 1987-04-29 | 1988-10-11 | Mobil Oil Corporation | Feedstock preparation and conversion of oxygenates to olefins |
| US4709113A (en) * | 1987-04-29 | 1987-11-24 | Mobil Oil Corporation | Conversion of crude methanol to gasoline with extraction |
| US4857667A (en) * | 1987-04-29 | 1989-08-15 | Mobil Oil Corporation | System for conversion of crude oxygenate to gasoline with feedstock extraction |
| US4827046A (en) * | 1988-04-11 | 1989-05-02 | Mobil Oil Corporation | Extraction of crude methanol and conversion of raffinate |
| US5009859A (en) * | 1988-04-11 | 1991-04-23 | Mobil Oil Corporation | Extraction and reactor system |
| US4827045A (en) * | 1988-04-11 | 1989-05-02 | Mobil Oil Corporation | Etherification of extracted crude methanol and conversion of raffinate |
| US4851606A (en) * | 1988-04-25 | 1989-07-25 | Mobil Oil Corporation | Control of waste water chemical oxygen demand in an oxygenate to hydrocarbon conversion process |
| US4981491A (en) * | 1989-07-28 | 1991-01-01 | Mobil Oil Corporation | Production of ether-rich fuel |
| US5233113A (en) * | 1991-02-15 | 1993-08-03 | Catalytica, Inc. | Process for converting lower alkanes to esters |
| US5306855A (en) * | 1991-02-15 | 1994-04-26 | Catalytica, Inc. | Catalytic process for converting lower alkanes to esters, alcohols, and to hydrocarbons |
| US5191142A (en) * | 1991-12-23 | 1993-03-02 | Amoco Corporation | Process for converting methanol to olefins or gasoline |
| US5750799A (en) * | 1995-03-15 | 1998-05-12 | Starchem, Inc. | Dimethyl ether production and recovery from methanol |
| US7102048B2 (en) * | 2002-12-17 | 2006-09-05 | Exxonmobil Chemical Patents Inc. | Methanol feed for producing olefin streams |
| KR100599251B1 (en) * | 2003-09-20 | 2006-07-13 | 에스케이 주식회사 | Catalysts for the dimethyl ether synthesis and its preparation process |
| US7879920B2 (en) | 2004-12-22 | 2011-02-01 | Exxonmobil Chemical Patents Inc. | Oxygenate to olefin manufacture and recovery process |
| DE102007022175B4 (en) * | 2007-05-11 | 2015-11-05 | Air Liquide Global E&C Solutions Germany Gmbh | Process and plant for the production of synthetic fuels |
| US8399728B2 (en) * | 2008-10-29 | 2013-03-19 | Lummus Technology Inc. | Absorber demethanizer for methanol to olefins process |
| US8445740B2 (en) * | 2008-10-29 | 2013-05-21 | Lummus Technology Inc. | Absorber demethanizer for FCC process |
| US20150073188A1 (en) * | 2012-03-01 | 2015-03-12 | The Trustees Of Princeton University | Processes for producing synthetic hydrocarbons from coal, biomass, and natural gas |
| DE102012112840A1 (en) | 2012-12-21 | 2014-06-26 | L'Air Liquide, Société Anonyme pour l'Etude et l'Exploitation des Procédés Georges Claude | Process and plant for the production of olefins |
| JP6362126B2 (en) * | 2013-03-11 | 2018-07-25 | 三菱重工エンジニアリング株式会社 | Methanol plant and gasoline synthesis plant |
| WO2018007484A1 (en) * | 2016-07-08 | 2018-01-11 | Haldor Topsøe A/S | Production of gasoline reformate blendstock |
| WO2018007487A1 (en) * | 2016-07-08 | 2018-01-11 | Haldor Topsøe A/S | Optimizing a mta process |
| CN110461997A (en) * | 2017-04-05 | 2019-11-15 | 埃克森美孚研究工程公司 | Integrated form separating methanol and preparing gasoline by methanol conversion process |
| WO2020150053A1 (en) * | 2019-01-18 | 2020-07-23 | Exxonmobil Research And Engineering Company | Conversion of methanol to gasoline with integrated paraffin conversion |
| CA3158380A1 (en) | 2019-10-24 | 2021-04-29 | Haldor Topsoe A/S | A process for the conversion of light alkanes to aromatic compounds with improved selectivity |
| EP4326833A1 (en) | 2021-04-20 | 2024-02-28 | Topsoe A/S | Process and plant for improving gasoline yield and octane number |
Family Cites Families (11)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US2591672A (en) * | 1949-01-03 | 1952-04-08 | Standard Oil Dev Co | Dehydration of alcohols by gasoline extractive distillation |
| US3293154A (en) * | 1963-02-15 | 1966-12-20 | Commercial Solvents Corp | Process for separating water from aqueous methanol mixtures |
| US3931349A (en) * | 1974-09-23 | 1976-01-06 | Mobil Oil Corporation | Conversion of methanol to gasoline components |
| US4014667A (en) * | 1975-06-16 | 1977-03-29 | Phillips Petroleum Company | Antifreeze recovery system |
| US4083888A (en) * | 1977-05-26 | 1978-04-11 | Mobil Oil Corporation | Process for manufacturing ethylene |
| US4349415A (en) * | 1979-09-28 | 1982-09-14 | Critical Fluid Systems, Inc. | Process for separating organic liquid solutes from their solvent mixtures |
| US4665249A (en) * | 1984-12-19 | 1987-05-12 | Mobil Oil Corporation | Method for feeding an MTG conversion reactor |
| JPS61280443A (en) * | 1985-06-05 | 1986-12-11 | Sumitomo Chem Co Ltd | Production of methyl tertiary butyl ether |
| US4709113A (en) * | 1987-04-29 | 1987-11-24 | Mobil Oil Corporation | Conversion of crude methanol to gasoline with extraction |
| AU613611B2 (en) * | 1987-04-29 | 1991-08-08 | Mobil Oil Corporation | Feedstock dewatering and etherification of crude methanol |
| US4777321A (en) * | 1987-04-29 | 1988-10-11 | Mobil Oil Corporation | Feedstock preparation and conversion of oxygenates to olefins |
-
1987
- 1987-04-29 US US07/043,718 patent/US4709113A/en not_active Expired - Fee Related
-
1988
- 1988-01-27 CA CA000557482A patent/CA1291770C/en not_active Expired - Lifetime
- 1988-02-01 AU AU11140/88A patent/AU609771B2/en not_active Ceased
- 1988-02-01 ZA ZA88693A patent/ZA88693B/en unknown
- 1988-02-02 NZ NZ223387A patent/NZ223387A/en unknown
- 1988-03-22 JP JP63065968A patent/JPS63275691A/en active Pending
- 1988-04-22 EP EP88303668A patent/EP0289234A1/en not_active Withdrawn
Also Published As
| Publication number | Publication date |
|---|---|
| EP0289234A1 (en) | 1988-11-02 |
| JPS63275691A (en) | 1988-11-14 |
| ZA88693B (en) | 1989-10-25 |
| AU609771B2 (en) | 1991-05-09 |
| US4709113A (en) | 1987-11-24 |
| NZ223387A (en) | 1989-12-21 |
| AU1114088A (en) | 1988-11-03 |
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