Production of diisopropyl ether

Abstract: In a diisopropyl ether production process the distillation step for the separation of aqueous IPA recovered from the DIPE extraction operations is modified to avoid carrying out the distillation under conditions that yield IPA-water azeotrope. The overhead and bottom streams from an aqueous IPA fractionation process operated under off-azeotrope conditions in DIPE production are recycled, respectively, to the DIPE reactor and extractor operations. Recycling the water and IPA overhead mixture eliminates the requirement for adding fresh water to the DIPE reactor. Returning the water and IPA bottom stream to the extractor reduces the requirement for distilled water addition to the extraction step. A less complex and costly aqueous IPA separation process results.

Description: This invention relates to processes for the production of diisopropyl ether (DIPE) and incorporates improvements in the product recovery operations for DIPE that enhance product separation while advantageously influencing overall etherification reaction operating conditions. Lower molecular weight alcohols and ethers such as isopropanol (IPA) and diisopropyl ether (DIPE) are in the gasoline boiling range and are known to have high blending octane numbers. In addition, by-product propylene from which IPA and DIPE can be made is usually available in a fuels refinery. An important aspect of research in the petroleum industry relates to processes to produce high octane lower aliphatic alkyl ethers as octane boosters and supplementary fuels. The catalytic hydration of olefins, particularly C₃ and C₄ olefins, to provide alcohols and ethers is a well-established art and is of significant commercial importance. Representative olefin hydration processes are disclosed in US-A-2,262,913, 2,477,380, 2,797,247, 3,798,097, 2,805,260, 2,830,090, 2,861,045, 2,891,999, 3,006,970, 3,198,752, 3,810,848 and 3,989,762. The production of ether from secondary alcohols such as isopropanol and light olefins is known. As disclosed in US-A-4,182,914, DIPE is produced from IPA and propylene in a series of operations employing a strongly acidic cation exchange resin as catalyst. Recently, processes for the hydration of olefins to provide alcohols and ethers using zeolite catalyst have been disclosed, for instance in US-A-4,214,107, 4,499,313, 4,757,664 and 4,906,187. In the conversion of a water feedstream and a C₃ hydrocarbons feedstream comprising propene and propane to DIPE and IPA as conventionally practiced, the conversion per pass is about 60%. The reaction effluent is a mixture containing unreacted water, C₃ hydrocarbons and hydrocarbon oligomeric by-products, in addition to DIPE and IPA. Separating these components requires multiple distillation and extraction operations that represent a substantial part of the overall process costs. C₃ and any lower hydrocarbons present are effectively removed by distillation. However, separation of DIPE and IPA is accomplished by an aqueous extraction operation that requires a further distillation step to separate alcohol and water. This extraction and distillation of IPA from the reaction effluent in order to recycle IPA to the etherification reactor as typically carried out, for instance as disclosed in EP-A-323,138, imposes an inordinate cost burden on the process in view of the formation of an IPA-water azeotrope which affects the complexity of the distillation tower design and operation. It is an object of the present invention to provide a process for the production of diisopropyl ether at lower overall process cost and complexity, in particular by improving product and IPA separation in downstream DIPE operations. It has been discovered that the DIPE process and distillation step for the separation of aqueous IPA recovered from the DIPE extraction operations can be modified and improved to avoid carrying out the distillation under conditions that yield IPA-water azeotrope and that such modifications beneficially affect overall process cost. We have found that the overhead and bottom streams from an aqueous IPA fractionation process operated under off-azeotrope conditions in DIPE production can be recycled, respectively, to the DIPE reactor and extractor operations. Recycling the water and IPA overhead mixture eliminates the requirement for adding fresh water to the DIPE reactor while returning the water and IPA bottom stream to the extractor reduces the requirement for distilled or treated water addition to the extraction step. As a consequence of these unforeseeable process advantages, a less complex and less costly aqueous IPA fractionation process can be implemented. According to the present invention a process for producing diisopropyl ether comprises contacting a feedstream containing propene and water under hydration and etherification conditions with an acidic etherification catalyst, separating from the effluent from said contacting a C₃⁻ stream and a bottom stream containing diisopropyl ether and isopropanol, extracting said bottom stream with water, and non-azeotropically separating the resulting aqueous extract phase to obtain a non-azeotropic aqueous isopropanol-rich stream and a water-rich stream, said isopropanol-rich stream being recycled to form part of said feedstream and said water-rich stream is used for extracting said bottom stream. The isopropanol-containing stream advantageously contains sufficient water to effect hydration of said propene, and is of water/isopropanol molar ratio greater than that of water-isopropanol azeotrope: typically it will provide for an isopropanol WHSV of about 0.10 and a water WHSV of about 0.05. The separation of the aqueous extract phase is preferably so carried out as to provide an isopropanol-containing stream at about 38°C at a pressure of about 70 kPa and a water-containing stream at a temperature of about 117°C. The etherification catalyst suitably comprises zeolite ZSM-5, zeolite Beta or an acidic resin, and the etherification conditions typically comprise a temperature of about 160°C and pressure of about 10,500 kPa. The source of feedstream propene may be a refinery C₃ hydrocarbon stream, additionally containing propane, and said bottom stream may additionally contain C₃+ olefins. In addition to an aqueous phase, the extracting of said bottom stream will yield an organic phase containing diisopropyl ether: advantageously the bottom stream is cooled before extraction. The separation of said effluent and/or of said aqueous extract phase is usually effected by distillation. In a specific embodiment of the invention, a process for the production of diisopropyl ether comprises contacting C₃ hydrocarbon feedstock containing propene and recycle stream from isopropanol separator overhead, said recycle stream comprising a non-azeotropic mixture of water and isopropanol, with acidic olefin hydration and etherification catalyst under etherification conditions at elevated pressure to produce an effluent stream containing diisopropyl ether, isopropanol, unreacted water, unreacted C₃ hydrocarbons and higher olefinic hydrocarbons; separating the effluent stream at reduced pressure to produce a separator overhead stream comprising C₃-hydrocarbons and a bottom stream containing diisopropyl ether, isopropanol, unreacted water and higher olefinic hydrocarbons; extracting the bottom stream and the water reactant feed, or fresh water make-up to the process, from separation of the reactor effluent stream in contact with a portion of aqueous bottom stream from the isopropanol separator to produce an extractor aqueous phase enriched in isopropanol and an organic phase comprising said diisopropyl ether and higher hydrocarbons; separating the extractor aqueous phase to provide an overhead stream comprising the non-azeotropic water and IPA recycle stream and the aqueous bottom stream which is recycled to the extraction operation. The single Figure is a process flow diagram of a preferred embodiment of the invention. In the DIPE process a feedstock comprising propene or a refinery C₃ hydrocarbon stream comprising olefins and paraffins, i.e., propene and propane, is contacted at elevated pressure with an acidic catalyst and water as a reactant to hydrate propene to form isopropanol (IPA) and etherify IPA to DIPE. Minor amounts of addition products of propene are also formed in the acidic catalyst environment, particularly hexenes and nonenes. On a per pass basis, the conversion of propene generally is about 60%, or between 50% and 70%. The effluent from the hydration and etherification zone is conventionally passed to a fractionator wherein a bottom stream is separated containing IPA and DIPE and an overhead stream that contains the unreacted C₃ hydrocarbons comprising propene and propane, if an olefin and paraffin feedstock has been used. The C₃ stream, typically containing both propene and propane, can be recompressed and recycled to the pressurized DIPE reactor. Since the recycle stream is rich in propane the DIPE feedstock stream pressure is increased in order to maintain the partial pressure of propene in the reactor. To avoid this, the recycle stream may be fractionated to purify propene recycle. Conventionally, DIPE is recovered by distillation and/or extraction of the fractionator bottom stream. This recovery system also separates an IPA stream and a water stream. The IPA stream is recycled to the etherification zone. The operating conditions of the olefin hydration and etherification process are not especially critical and include a temperature of from about 60 to 450°C, preferably from about 90 to about 220°C and most preferably from about 120 to about 200°C, a pressure of from about 100 (700 kPa) to about 3500 psi (24,500 kPa), preferably from about 500 (3500 kPa) to about 2000 psi (14,000 kPa), a water to olefin mole ratio of from about 0.1 to about 30, preferably from about 0.2 to about 15 and most preferably from about 0.3 to about 3. The olefin hydration process can be carried out under dense phase, liquid phase, vapor phase or mixed vapor-liquid phase conditions in batch or continuous manner, preferably using a fixed bed reactor. A liquid hourly space velocity (LHSV) of from about 0.1 to about 20, preferably about 0.1-2, when operating continuously is suitable. The catalyst employed in the olefin hydration and etherification operations may be an acidic resin catalyst such as sulfonated polystyrene Alternatively a shape-selective acidic zeolite catalyst can be used, including the zeolites Y, Beta, ZSM-35 and MCM-22 (described in US Patent Application Serial No. 456,702, filed December 26, 1989). The invention may be put into practice by the use of a reactor system comprising a combination of reactor means for hydration and etherification of C₃ hydrocarbons; first distillation means operatively connected to said reactor means for separation of effluent therefrom; extractor means receivably connected to a bottom portion of said first distillation means for aqueous extraction of distillation effluent and having inlet means for receiving aqueous extraction media and outlet means for discharging aqueous extract and organic phases; second distillation means operatively connected to said extractor means for separating aqueous extract therefrom; first conduit means, operatively connected to said second distillation means and extractor means, for passing bottom stream from said second distillation means to said extractor means; and second conduit means, operatively connected to said second distillation means and said reactor means, for passing overhead stream from said second distillation means to said reactor means. The following Example describes a process in accordance with the invention, and illustrates the novel downstream separation of the products of the hydration and etherification reaction and, more particularly, to the operation of the IPA-water fractionator with recycle streams provided to the extractor and reactor section. Example Referring to the Figure and to Table 1, a feed 110 containing about 157 parts of propene per 100 parts of propane is washed with water to remove basic nitrogen components and pumped 112 in combination with an isopropanol/water recycle stream 115 to a DIPE reactor 114 operated at a pressure of about 1500 psig (10,500 kPa) and 320°F (160°C). Typically, the feedstream comprises FCC propane/propene and is fed to parallel DIPE tubular reactors. The propene reacts with water over a catalyst comprising zeolite beta to produce DIPE at a conversion of about 57%. The reaction products exit the reactor as an effluent 116 at a temperature of about 340°F (171°C) and are cooled 118 to about 150°F (66°C). The effluent contains about 100 parts of DIPE and 63 parts of IPA per 100 parts of propane feed as shown in Table 1. The cool effluent is passed via expander 120, to reduce the pressure to about 250 psig (121 kPa) and a temperature of about 174°F (79°C), to a 15-tray propane/propene splitter 122 whence C₃ and lighter material are removed as an overhead stream 124 at a temperature of about 100°F (38°C). The bottoms stream 126, exiting the splitter at a temperature of about 275°F (135°C) and a pressure of 60 psig (420 kPa), is let down in pressure in a valve and cooled to about 100°F (38°C) by exchanging heat against an extractor bottoms stream 128 and cooling water. The cooled stream 130 is then fed to an extractor 132 which operates at a pressure of 40 psig (280 kPa) and water:feed volume ratio of 3:1. Water is fed to extractor 132 as a recycle stream 134 from IPA column 144 following pumping 136 and heat exchange to provide a recycle inlet stream temperature to the extractor of about 100°F (38°C). A make-up boiler feed water stream 138 is also added to the extractor to compensate for any water losses. The ether phase, containing 89% DIPE, 9.4% C₅=+, 1% water and traces of C₄s and IPA is removed as extractor overhead 140 at a temperature of about 100°F (38°C) and 40 psig (280 kPa). The aqueous extractor bottom stream 128 is preheated in succession to about 180°F (82°C) against splitter bottoms as previously noted and IPA column bottoms 142 before being fed to IPA column 144. The IPA column is a 15-tray column which operates at a pressure of about 10 psig (70 kPa). The overhead stream 146 from the column is recycled back to the reactor through pump 148 at a temperature of about 100°F (38°C) and 1600 psig (11,200 kPa) to provide stream 115 at the required IPA weight hourly space velocity (WHSV) of 0.10 and water WHSV of 0.05. The IPA column bottoms stream at a temperature of about 117°C is recycled to the extractor after a small slip stream 150 is taken to maintain high purity. Table 1 shows that the recycle overhead stream from the IPA column contains an amount of water considerably above that contained in IPA-water azeotrope. Consequently, the fractionator design is less complex than that conventionally used in DIPE production where the IPA column is designed and operated to produce an overhead containing a minimal amount of water. The discovery of the utility of the water rich IPA overhead stream and IPA rich aqueous bottom stream as described herein permits the design and operation of a much simpler and less costly IPA column.

Claims: A process for producing diisopropyl ether comprising contacting a feedstream containing propene and water under hydration and etherification conditions with an acidic etherification catalyst and separating from the effluent from said contacting a C₃⁻ stream and a bottom stream containing diisopropyl ether and isopropanol, characterised by extracting said bottom stream with water, and non azeotropically separating the resulting aqueous extract phase to obtain a non-azeotropic aqueous isopropanol-rich stream and a water-rich stream, said isopropanol-rich stream being recycled to form part of said feedstream and said water-rich stream being used for extracting said bottom stream. A process according to claim 1 wherein said isopropanol-rich stream contains sufficient water to effect hydration of said propene. A process according to any preceding claim wherein said recycled stream provides for an isopropanol WHSV of about 0.10 and a water WHSV of about 0.05. A process according to any preceding claim wherein the separation of said aqueous extract phase is so carried out as to provide an isopropanol-rich stream at about 38°C at a pressure of about 70 kPa and a water-rich stream at a temperature of about 117°C. A process according to any preceding claim wherein said etherification catalyst comprises zeolite ZSM-5, zeolite Beta or an acidic resin. A process according to any preceding claim wherein said etherification conditions comprise a temperature of about 160°C and pressure of about 10,500 kPa. A process according to any preceding claim wherein the source of said propene is a refinery C₃ hydrocarbon stream and said feedstream additionally contains propane. A process according to any preceding claim wherein said bottom stream additionally contains C₃+ olefins. A process according to any preceding claim wherein, in addition to said aqueous phase, the extracting of said bottom stream yields an organic phase containing diisopropyl ether. A process according to any preceding claim wherein said bottom stream is cooled before extraction. A process according to any preceding claim wherein the separation of said effluent and/or of said aqueous extract phase is effected by distillation.

Inventors: HARANDI MOHSEN NADIMI, OWEN HARTLEY NMN, RAGONESE FRANCIS PAUL, STOOS JAMES ARTHUR, HARANDI, MOHSEN NADIMI, OWEN, HARTLEY, (NMN), RAGONESE, FRANCIS PAUL, STOOS, JAMES ARTHUR
Applicants: MOBIL OIL CORP, MOBIL OIL CORPORATION
Application year: 1991
Main classification: C07C41
IPC codes: B01J31, C07B61, C07C43, C07C41, B01J29