CS264304B2 - A method for producing dimethyl ether, optionally together with methanol - Google Patents

Abstract

A process for the preparation of dimethyl ether, optionally together with methanol, by conversion of a synthesis gas in which the synthesis gas with an Hj / CO molar ratio in the range of 0.5 to 3.0 is brought into contact, optionally in the presence of water, at 232 to 399 ° C. with a catalyst comprising co-precipitated Cr, Cu and Zn components and an acidic dehydrating component selected from the group consisting of gamma-alumina, silica-alumina, crystalline zeolite ZSM-5 with high silica content, phosphate, titanium dioxide in combination with silica, rare soil and clay in which the Cr / (Cu + Zn) atomic ratio is 0.1 to 1.0 and the acid dehydration content is 5 to 95% by weight, whereupon the catalyst is regenerated by contacting the oxygen-containing gas at a temperature of 38 to 538 ° C.

Description

The invention relates to a method for converting synthesis gas into dimethyl ether.

Two main processes are currently available for converting coal to liquid fuels via synthesis gas, the well-known and published Fischer-Tropsch (FT) process and the more recently developed methanol-gasoline process, such as that described in OSA Patent No. 3,928,483. The Fischer-Tropsch process produces a wide range of products up to _C5Q , including gases, liquid hydrocarbons, oxygenates and water.

Common to both processes is the overwhelming influence of the investment costs in producing synthesis gas, i.e. a mixture of hydrogen and carbon monoxide. The costs vary with the type of gasifier and the choice of gasifier is influenced by the properties of the coal and the enthalpy of the product. A recent review of gasification technology found some highly efficient gasifiers which use small amounts of steam and oxygen in the gasification and therefore necessarily produce synthesis gas with a very low hydrogen to carbon monoxide ratio. The most recent advanced gasifiers operate in such a way that the product is obtained at relatively low temperatures, at a low steam to carbon ratio, thereby achieving high thermal efficiency. Synthesis gas produced under these conditions may have a molar ratio of hydrogen to carbon monoxide equal to one and more usually in the most efficient gasifiers this ratio is less than 1, in the range of about 0.4 to about 0.7. Syngas with such a low H2/CO molar ratio cannot be directly used in current conventional FT processes or in methanol synthesis. Both of these processes require an _H2 /CO ratio of 2 or higher. However, any external conversion of the low _H2 /CO molar ratio syngas to water gas, in order to increase the _H2 /CO molar ratio from less than 1 to 2 or higher, would essentially eliminate any coal-to-syngas conversion efficiency advantages achieved using the most advanced high-efficiency gasifiers. '

The main advantage of producing dimethyl ether (DME) directly from synthesis gas is that this compound can be easily converted to hydrocarbons with a gasoline boiling range using crystalline zeolites that have a steric hindrance index (C-index, constraint index) of 1 to 12, see for example U.S. Patent No. 3,928,483. In this process, the methanol formed is dehydrated and the ether product is converted on a crystalline zeolite of the ZSM-5 type.

One problem encountered in the synthesis of methanol from synthesis gas is that it achieves a relatively low equilibrium conversion (efficiency on carbon). In order to achieve satisfactory yields, large amounts of unreacted hydrogen and carbon monoxide must be recycled in conventional methanol synthesis. This factor significantly affects the economic attractiveness of this method, since the need to recycle large amounts of substances requires a significant increase in the size of the plant, which increases investment costs.

Consequently, improvements were sought to increase the methanol productivity per pass through the reaction zone in the conversion of syngas. One approach was to remove the chemical equilibrium by further reacting the methanol to a non-equilibrium final product. Among the reactions that removed the reaction equilibrium was the dehydration of the intermediate methanol product to dimethyl ether.

For example, U.S. Patent No. 3,894,102 describes a two-stage process for converting synthesis gas to gasoline, in which, in the first stage, the synthesis gas is contacted with a mixed catalyst comprising a methanol synthesis catalyst component and an acidic dehydration component to form a dimethyl ether-containing product, and then, in the second stage, this dimethyl ether-containing intermediate product is contacted with a ZSM-5 catalyst to form aromatic hydrocarbons. Any catalyst for methanol synthesis can be used in the first stage, although a mixture of oxides of the composition Zn/Cu/Cr/rare earth metals is preferably used, in which the content of the individual components, expressed in parts by weight of metal, is 50 to 70 parts copper, 50 to 25 parts zinc, 5 to 15 parts chromium and 5 to 15 parts rare earth metals. The Cr/(Cu+Zn) ratio in these mixtures is in the range of 0.07 to 0.3.

The synthesis gas used in the process of U.S. Patent No. 3,894,102 had a relatively high molar ratio of H ₂ /CO. Another important improvement that can be achieved in connection with the reaction to dimethyl ether, which reaction removes the equilibrium in the conversion of synthesis gas to methanol in the presence of catalysts for the synthesis of methanol, is disclosed in U.S. Patent No. 4,011,275. In the process of this patent, a synthesis gas with a molar ratio of H ₂ /CO of 0.8 to 1.7 is used for the conversion, which gas is passed over a two-component catalyst for the synthesis of dimethyl ether. The conversion of the synthesis gas in this synthesis of dimethyl ether and methanol is higher than that in the case of the synthesis of methanol alone, but not only that, the maximum equilibrium conversions in the case of the synthesis of dimethyl ether and methanol are achieved at a molar ratio of H ₂ /CO in the synthesis gas of 1, in contrast to the value of 2, which is optimal for the synthesis of methanol alone. Figures 6 and 8 of the latter patent further show that the addition of water to the synthesis gas feed further shifts the _H2 /CO ratio required for maximum equilibrium conversion towards feeds leaner in hydrogen. The addition of carbon dioxide to the feed had no apparent effect on the synthesis of dimethyl ether and methanol. The increase in conversion achieved using hydrogen-lean feeds in the case of simultaneously introduced water was attributed to an internal reaction to water gas between the added water and the excess carbon monoxide. This reaction produces additional hydrogen which can be converted to an organic product. It is stated that all conventional catalysts are suitable as catalysts for the synthesis of methanol in this process.

U.S. Patent No. 4,098,809 discloses that dimethyl ether can be produced from a hydrogen-poor feed mixture containing carbon monoxide, carbon dioxide and hydrogen by contacting the mixture under conversion conditions with a Cu/Zn/Cr/alumina methanol synthesis catalyst, provided that the carbon dioxide content of the feed gas is relatively low relative to the carbon monoxide content. The patent discloses that the presence of water in the reacting gas is undesirable because it inhibits the dehydration of methanol and thereby limits the degree of conversion to dimethyl ether achieved in a single pass through the reaction zone. However, it is noted that the addition of steam is useful when the carbon monoxide is in substantial excess relative to hydrogen. Suitable catalytic components for methanol synthesis in this process are copper, zinc and chromium catalysts. In the examples, a catalytic component for methanol synthesis with a composition of Cu/Zn/Cr in an atomic ratio of 82/16/4 was used. The Cr/(Cu+Zn) ratio in this catalyst was 0.04. The patent does not disclose any methods for preparing the catalyst.

U.S. Patent No. 4,096,163 proposes converting synthesis gas to hydrocarbon products over a catalyst comprising a methanol synthesis component and a ZSM-5 type zeolite component.

There is no indication that any catalyst for methanol synthesis is particularly preferred.

In addition, Sherwin and Blum reported in the May 1978 Interim Report, "Liquid Phase Methanol," prepared for the Electric Power Research Institute, Palo Alto, California, an attempt to produce dimethyl ether by adding gamma-alumina and 13X molecular sieves to a slurry reaction system containing a commercially available methanol synthesis catalyst. At temperatures of 230 to 300°C, absolute pressures of 3.5 to 6.9 MPa, and gas hourly space velocities of 2,015 to 6,915 h ^-1, only traces of dimethyl ether were observed. A catalyst containing copper, zinc and chromium in the atomic ratio Cu/Zn/Cr 6/3/1 impregnated on Davidson 980 silica-alumina provided only trace amounts of dimethyl ether at a temperature of 230 °C, an absolute pressure of 6.9 MPa and a gas hourly space velocity of 2,000 h ^-1 when using a feed containing 50% hydrogen, 25% carbon monoxide, 10% carbon dioxide and 15% methane.

U.S. Patent No. 4,177,167 discloses that catalysts for the synthesis of dimethyl ether which have been stabilized with silicon compounds have higher activity and longer service life than unstabilized catalysts. The catalysts consist of mixtures of oxides or salts of chromium, lanthanum, manganese, copper, zinc, and combinations of those oxides or salts with aluminum. To maintain the activity of the stabilized catalyst, an atomic aluminum content in the range of 10 to 70% is required. The catalyst can be prepared by any conventional process. The patent exemplifies the coprecipitation of Cu/Zn/Cr components. The composition of the catalyst (other than the aluminum content) is not stated to have any particular effect. The Cr/(Cu+Zn) ratio in the catalysts given by way of example is in the range of approximately 0.2 to 0.6. The patent also states that a suitable molar ratio of H2/CO in the feed gas is in the range of about 0.3 to about 10. The synthesis gas used in the examples had a molar ratio of _H2 /CO3.

The feasibility of direct synthesis of dimethyl ether was initially demonstrated using a physical mixture of Cu/Zn/Cr containing a rare earth metal-based promoter and gamma-alumina,

Synthesis gas with a molar ratio of H-/CO = 4 was contacted with the catalyst at 315 °C, an absolute pressure ^of 5 MPa, and a space hourly gas velocity of 1,440 h -1 . A conversion of 31.5% (86% equilibrium conversion) was achieved. The product contained 22.4% dimethyl ether, 0.93% methanol, 29.0% carbon dioxide, 2.0% water, 3.2% light hydrocarbons, and the remainder was hydrogen and carbon monoxide. The above catalyst was also tested at a molar ratio of H ₂ /CO = 1, a temperature of 315 °C, an absolute pressure of 10.3 MPa, and a space hourly gas velocity of 4,400 h ^-1 . The conversion dropped from 60% (75% equilibrium conversion) to 40% within 30 days, indicating that the catalyst was aging rapidly. An attempt at reactivation with hydrogen at 315 to 40 °C was not successful.

Catalyst aging is a significant problem in known dimethyl ether synthesis processes, as aging severely limits the commercial attractiveness of otherwise highly valuable processes.

The aim of the invention was to find a reliable and commercially attractive process for regenerating catalysts that have aged by being deactivated by use in the synthesis of dimethyl ether. The ultimate goal was to increase the economic attractiveness of the entire synthesis process.

The invention is based on the discovery that catalyst aging can be minimized by using a certain Zn/Cu/Cr catalyst as a component of a dimethyl ether synthesis catalyst that catalyzes methanol synthesis, and that the catalyst activity can be maintained at a high level sufficient to ensure near-equilibrium conversion of the feed synthesis gas by performing regeneration by a special oxidation technology.

Although the use of such catalysts in the synthesis of dimethyl ether was already known, the relationship between the composition of the catalyst, in particular its sulfur content, and the aging of the catalyst had not previously been clearly appreciated.

The subject of the invention is a method for producing dimethyl ether, optionally together with methanol, by converting synthesis gas, which is characterized in that synthesis gas with a molar ratio of H ₂ /CO in the range from 0.5 to 3.0 is brought, optionally in the presence of water, at a temperature of 232 to 399 °C into contact with a catalyst containing co-precipitated Cr, Cu and Zn components and an acidic dehydration component selected from the group comprising gamma-alumina, silica-alumina, crystalline zeolite of the ZSM-5 type with a high silica content, phosphate, titanium dioxide in combination with silica, rare earth and clay, in which the atomic ratio Cr/(Cu+Zn) is 0.1 to 1.0 and the content of the acidic dehydration component is 5 to 95% by weight, after which the catalyst is regenerated by bringing into contact with a gas containing oxygen at a temperature of 38 to 538 °C.

In the synthesis gas conversion according to the invention, metal components of the catalyst for methanol synthesis are used in combination with an acidic dehydration component. The catalysts used in the process according to the invention are prepared in particular by the technology of co-precipitation of the hydrogenation components either alone or mixed with the dehydration component, the atomic ratio of Cu, Zn and Cr varying within the relatively narrow limits specified above. The co-precipitated components Cu, Zn and Cr are used in such relative amounts that the ratio Cr/(Cu+Zn) is preferably 0.25 to 0.75. The ratio 0.5 is particularly preferred. The ratio Cu/Zn is preferably 0.5 to 3.0.

Of the acidic dehydration components mentioned above, gamma-alumina is particularly preferred, which is preferably used in an amount corresponding to about 50% by weight of the catalyst.

The process for the synthesis of dimethyl ether according to the invention is of particular interest because the molar ratio _H2 /CO can be lower than 1 or higher than 1. The molar ratio _H2 /CO can therefore be in the range from 0.5 to 3.0, preferably from 0.5 to 2. However, it is particularly preferred to use a synthesis gas with a molar ratio of 1 or lower, since gas with such a ratio is produced more economically in modern gasifiers with high efficiency, as mentioned above. The use of such a synthesis gas source can result in a 30 to 40% reduction in operating costs. It has been found that when using a synthesis gas with a molar ratio H2/CO below 1, for example in the range from 0.4 to 0.7, the hydrogen deficiency can be corrected or compensated for by injecting water vapor into the catalyst, either separately or in a mixture with the low H2/CO ratio feed synthesis gas. The added steam reacts with the specific catalyst used to form water gas, thereby effectively increasing the _H2 /CO ratio in the gas contained in the reaction zone containing the catalyst. In this way, it is possible to avoid the need to perform an external conversion of the synthesis gas to water gas in order to increase the molar ratio of _H2 /CO in the synthesis gas used. This process further contributes to the overall economic improvement of the process.

Research conducted in the development of a process for converting synthesis gas to dimethyl ether and further studies following on from this research have shown that catalysts of a certain composition consisting of coprecipitated oxides of copper, zinc and iron mixed with a suitable acidic matrix such as gamma-alumina provide high catalytic activity and reactivity in carrying out the synthesis of dimethyl ether. More importantly, however, it has been found that catalysts of this composition can be periodically oxidatively regenerated to maintain their high catalytic activity. Thus, the catalysts of the composition used in accordance with the invention can be maintained at high and constant activity over a long period of operation by carrying out a certain oxidative regeneration and pretreatment of relatively short duration. For example, the catalyst in the synthesis gas reaction zone can be regenerated directly in this reaction zone or can be passed through a separate catalyst regeneration zone and then returned to the reaction zone with regenerated activity and selectivity. Catalyst regeneration can therefore be carried out continuously, semi-continuously or intermittently, for example periodically, depending on whether a fixed bed or a moving catalyst system is used.

It has now been found that, for example, the co-precipitated Cr, Cu, Zn components of the catalyst mixed with gamma-alumina can be maintained in a high state of activity to achieve substantially thermodynamic yields of dimethyl ether by periodically regenerating the catalyst with an oxygen-containing gas in which the partial pressure of oxygen is the same as in air or lower for a relatively short contact time or until the said oxygen-rich gas has penetrated the system. There is no evidence that this procedure has been previously carried out.

Using the coprecipitated catalyst described above, it has been found that the catalyst can be successfully regenerated at temperatures of 38 to 538°C, with temperatures below 538°C being preferred. After oxidation or regeneration of the catalyst, it is important, or rather crucial, to avoid contact of the freshly oxidized catalyst with a high temperature reducing gas.

That is, the reduction of the freshly regenerated catalyst with hydrogen, a mixture of hydrogen and carbon monoxide, or another reducing gas, should be avoided, especially when the contact temperature is 260°C or higher. It is therefore preferable to contact the oxygen-regenerated catalyst with the reducing gas or dilute reducing gas initially at a temperature of 177°C to about 232°C and then increase the temperature and/or change the composition of the synthesis gas used to achieve near thermodynamic yields of dimethyl ether.

Catalyst regeneration can be carried out periodically or continuously depending on whether a fixed or moving bed of catalyst is used. In addition, the conversion of synthesis gas with a low _H2 /CO ratio to dimethyl ether can be carried out under gradually milder conversion conditions, milder than the equilibrium conditions, in which case the catalyst can be kept in operation for a longer time than the oxidative regeneration is required. In the oxidative regeneration of the catalyst, excessively high temperatures resulting from exothermic reactions should be avoided. The catalyst particles being oxidatively regenerated should not come into contact with such temperatures. It is advantageous to use frequent catalyst regeneration, such as used in reaction systems with alternating reactors, where regeneration can be carried out after a week or even daily, or regeneration with a medium frequency. Daily regeneration of the catalyst mass is desirable when the synthesis gas is to be converted to dimethyl ether at substantially equilibrium yields. The regeneration temperature should preferably be maintained below about 538°C to achieve optimal catalytic efficiency.

The invention is described in more detail and illustrated below using the figures.

Figures 1 to 4 illustrate various aspects of the co-dosing of water and synthesis gas onto the catalyst.

Figure 5 shows a block diagram of the conversion of synthesis gas to liquid fuels.

Figure 6 illustrates the effect of catalyst composition on its aging.

Figure 7 illustrates the effect of catalyst preparation on its in-service properties.

Figures 8 and 9 illustrate individual aspects of catalyst regeneration.

Figures 10 to 14 illustrate individual aspects of catalyst preparation and the influence of catalyst composition on its properties during operation.

As already mentioned, the synthesis of dimethyl ether has decisive advantages compared to the Fischer-Tropsch process or methanol synthesis, since it can utilize synthesis gas with a relatively low H ₂ /CO ratio. As a result, this synthesis can be carried out advantageously in combination with a highly efficient gasification process, which produces synthesis gas with a relatively low H ₂ /CO molar ratio (less than 1). The lack of hydrogen is compensated by injecting or adding the necessary amount of water vapor together with the synthesis gas with a low H ₂ /CO molar ratio. In this case, the ability of the catalyst to catalyze the conversion of synthesis gas into water gas is used.

Fig. 1 shows the dependence of the synthesis gas conversion (H^+CO) in i molar (y axis) on the amount of water added in moles per 100 moles of synthesis gas. The graph illustrates one effect of adding water vapor. The molar ratio H^/CO was 0.67, the temperature 316 °C and the pressure 5, 7.5 and

MPa. As is evident from Fig. 1, the conversion passes through a maximum that shifts to a higher relative amount of steam fed with increasing pressure. It is therefore evident that the maximum conversions are somewhat lower than those achieved at comparable total pressures at a molar ratio of H ₂ /CO equal to 1.

Direct synthesis of dimethyl ether from synthesis gas can be used with considerable advantage in many different combined processes. The economical formation of dimethyl ether from synthesis gas with a low molar ratio of H ₂ /CO obtained from an efficient source or produced by in situ coal gasification is particularly suitable for competitive processes for the production of specific chemicals and/or hydrocarbons with widely varying compositions.

For example, the co-dosing of water with a synthesis gas with a low _H2 /CO ratio has a significant effect on increasing the utilization of the synthesis gas by converting excess carbon monoxide to water gas to produce additional hydrogen. An additional but equally important effect of the co-dosing of water is to increase the _H2 /CO ratio in the regenerated stream leaving the synthesis gas conversion to a value of at least 1. This effect of increasing the molar ratio of _H2 /CO in the regenerated outgoing stream is illustrated in Fig. 2, which is a graphical representation of the dependence of the molar ratio of _H2 /CO in the gas leaving the catalytic zone for the conversion of synthesis gas on the molar amount of water (steam) supplied to the catalyst per 100 moles of synthesis gas. The initial _H2 /CO ratio was 0.67, the temperature was 316 °C and the pressure was 5, 7.5 and 10 MPa.

In the synthesis according to the invention, the dimethyl ether product produced can be refined by reaction with a suitable alcohol, for example tert-butanol or a branched olefin, for example 2-methylpropene, to other ethers, for example methyl tert-butyl ether. An acidic catalyst is used in this case. Sulfuric acid or other inorganic acids can be used as the acidic catalyst, optionally with the use of a carrier. Alternatively, acidic ion exchange resins or phosphoric acid on a carrier can be used as the acidic catalyst. The reaction temperature can be from 50 to 300 °C, with a particularly preferred temperature range being 60 to 150 °C. A suitable pressure for this reaction is 0.3 to 5 MPa.

According to another method, the synthesis gas originally produced by gasification can be converted to a dimethyl ether-rich product in a catalyst slurry reaction system. In such a system, the metal and acid components of the catalyst are surrounded by a suitable high-boiling liquid phase by being suspended therein. For example, mineral oil, an oily product from the Fischer-Ropsch synthesis, a portion of a decanted oily product, and a base stock for lubricating oils can be used as liquid phases. The solid catalyst components can be used in the form of combined particles or in the form of separate particles containing the metal and acid components. The liquid phase reaction is typically carried out at a temperature of from about 260 to about 343 °C at a pressure of from about 6.8 to 13.6 MPa.

In a further process, the dimethyl ether product produced can be converted to methanol by reaction with water in the presence of an acidic catalyst. The catalyst can be phosphorus pentoxide on diatomaceous earth, silica-alumina, reduced tungsten oxide or acidic crystalline zeolite. The temperature is typically from about 204 to about 371°C and the pressure is from 1.0 to 10.0 MPa.

Alternatively, the dimethyl ether product produced can be reacted with oxygen in the presence of a copper catalyst to form formaldehyde. The temperature for this reaction can be from about 600 to about 720°C and the pressure from 0.1 to 0.5 MPa. Instead of the copper catalyst, tungsten oxide can be used in combination with phosphoric acid distributed on a carrier such as alumina, silica-alumina or diatomaceous earth, or silver alone or promoted with small amounts of copper and silver. When using a silver catalyst, the temperature is usually in the range of from 600 to 800°C and the pressure from 0.3 to 0.8 MPa. Furthermore, catalytic mixtures containing iron promoted by molybdenum (for example, with a mass composition of 18% Fe ₂ °3 ^{+ 82} % Moop at a temperature of 350 to 400°C) can also be used.

Liquid fuels can be prepared from the olefinic intermediate product of the dimethyl ether conversion by passing this stream to a conversion reaction for the production of liquid fuels. In this case, the dimethyl ether product is converted at a temperature of 343 to 427 °C, preferably at a temperature of at least 371 °C, to an olefinic intermediate product by contacting it with a crystalline zeolite of the class represented by the crystalline zeolite HZSM-5. The olefins formed may be branched or straight chain depending on the conditions used and the activity of the catalyst and conveniently have a boiling point in the range of the boiling point of gasoline. Alternatively, the operating conditions and catalyst can be selected such that the resulting olefins produce a gasoline rich in aromatics. The olefinic intermediate can be further upgraded to jet fuel by a combined oligomerization or olefination and hydrogenation process. The oligomerization of the intermediate olefin product can be carried out at a temperature of from 175 to about 250°C, preferably at a temperature of 200 to 220°C, under a pressure of 1.0 to 10 MPa. When oligomerizing olefins for jet fuel, it is convenient to also introduce water into the reaction zone and maintain about 15 to 18% phosphorus pentoxide in the catalyst. Diatomaceous earth can be used as an acidic support for the catalyst.

The synthesis of olefins from the dimethyl ether product is an intermediate step carried out in combination with steps leading to the production of chemical components, branched olefins, aromatics and substances with a boiling range corresponding to gasoline depending on the reaction conditions and the catalyst used. In addition to olefins produced from dimethyl ether, olefin mixtures from any other source can also be used for the reaction with methanol to produce high-octane ethers.

According to yet another alternative, unreacted synthesis gas can be separated from the dimethyl ether product. This unreacted synthesis gas, which normally has an I/CO ratio of at least 1, contains a significant amount of carbon dioxide. Synthesis gas of such composition can be advantageously used in the Fischer-Tropsch conversion of synthesis gas using iron or cobalt or ruthenium as the metal component of the Fischer-Tropsch catalyst. The unreacted synthesis gas from the dimethyl ether synthesis can therefore be subjected to the Fischer-Tropsch conversion process, optionally in the presence of a metal component catalyzing the conversion to water gas.

In each of the above processes, the unreacted synthesis gas can be recycled to the dimethyl ether conversion, after optionally separating the carbon dioxide and adding water to the recycled gas. Any unreacted gas containing carbon dioxide can, on the other hand, be fed to the Fischer-Tropsch synthesis gas conversion, with optionally adding water to the feed. This reaction yields a wide range of products.

Fig. 5 shows a block diagram of a unit for upgrading synthesis gas by treatment according to the invention and further treatment. Synthesis gas with a low molar ratio of Hj/CO, i.e. with a ratio of _H2 /CO of 1 or less, is introduced into the process through line 2^ to be converted into dimethyl ether in zone i. Water is introduced separately into zone £ through line J5 or water is introduced through line j1, mixed with synthesis gas in line 2 before being introduced into zone j1 for the synthesis of dimethyl ether.

In zone £, the temperature, pressure and space velocity conditions are selected to produce a product rich in dimethyl ether but containing some methanol. Zone _4 also contains a separation device by which unreacted synthesis gas and carbon dioxide are separated from the dimethyl ether-rich product. The unreacted synthesis gas and carbon dioxide are recovered in line 10 and are fed to further processing. The dimethyl ether-rich product containing methanol and carbon dioxide is discharged through line 12. Through line 12, the dimethyl ether-rich product is fed to zone 14 where the ether is converted to olefins with a boiling range corresponding to gasoline and/or aromatics by contact with a crystalline zeolite catalyst represented by the ZSM-5 catalyst. Also, the methanol contained in the obtained dimethyl ether product is converted to substances with a boiling range corresponding to gasoline upon contact with the zeolite catalyst. It is believed that olefination, polymerization and aromatization occur in zone 14 under operating conditions that are in each case selective with respect to the purpose of the catalyst used. The zeolitic catalyst component can be given some hydrogenation activity by adding one or more elements from the group consisting of cobalt, iron, platinum, palladium, ruthenium, rhodium, osmium, copper, zinc, manganese and molybdenum. As a product of this operation, an olefinic and/or aromatic product with a gasoline boiling range is obtained in line 16.

Synthesis gas with an increased molar ratio of H ₂ /CO containing carbon dioxide obtained in zone £ and discharged through line 10 is discharged to zone 18 containing the Fischer-Tropsch catalyst.

In this zone, the enriched synthesis gas is converted, optionally in the presence of added water.

Fischer-Tropsch type. In zone 18, Fischer-Tropsch synthesis reactions catalyzed by iron, cobalt and ruthenium, as catalysts, although other components such as osmium and rhodium can be used with varying success. The Fischer-Tropsch metal component may be mixed with a zeolite or with an acidic carrier modified to adjust the activity of the individual components to the desired range. The Fischer-Tropsch component may also be in admixture with a crystalline zeolite represented by crystalline zeolite ZSM-5 to effect the conversion of the synthesis gas to components having the boiling range of gasoline. The catalyst used in zone 18 may be in a fluidized bed or in a fixed bed or may be suspended in a liquid phase reaction system. In this manner, the synthesis gas may be converted over a Fischer-Tropsch catalyst, such as potassium-modified iron, to a mixture of up to carbonaceous components containing alcohols, oxygenates and hydrocarbons. Fischer-Tropsch synthesis is normally carried out at a temperature

149 to 427 °C and at a pressure of 0.35 to 13.6 MPa. The products of Fischer-Tropsch synthesis are normally separated into a low-boiling component, a gasoline-boiling range product, and a higher-boiling decanter oil. The oxygenates are recovered in the aqueous phase and then converted to hydrocarbons using ZSM-5 zeolites.

Upgrading of the Fischer-Tropsch synthesis product can be carried out in a number of different ways, as described in recent publications. Particularly suitable is the processing of the Fischer-Tropsch product into LPG (a mixture of propane and butane), gasoline and diesel fuel.

Upgrading of unreacted synthesis gas obtained from dimethyl ether can be carried out, for example, by the technique described in U.S. Patents Nos. 4,086,262, 4,157,338 and 4,159,995 and/or the entire Fischer-Tropsch synthesis product can be processed by the technique described in U.S. Patents Nos. 4,041,094, 4,046,830 and 4,046,831. Alternatively, the Fischer-Tropsch synthesis product can be separated so that the gasoline boiling point and the lower boiling point are processed by the technique of U.S. Patents Nos. 4,041,095, 4,095,505, 4,049,741 and 4,052,477. The higher boiling point than gasoline, referred to as decant oil, can be processed by the technique described in U.S. Patents Nos. 4,041,095, 4,095,505, 4,049,741 and 4,052,477. 044,064, 4,046,829, 4,053,532, 4,059,648 and 4,071,574.

In the process illustrated in FIG. 5, in accordance with a specific embodiment of the process, a synthesis gas containing carbon dioxide is processed over a Fischer-Tropsch catalyst in zone Π3 to a product containing alcohols, oxygenates and hydrocarbons. The product is separated into hydrocarbons in zone 18 which are passed through line 20 to zone 22. In zone 22, the hydrocarbons are contacted with one or more catalyst beds containing crystalline zeolite of the ZSM-5 type having pore openings of at least 0.5 nm, a silica to alumina ratio of greater than 12 and a barrier index in the range of 1 to 12. In this catalytic process, the temperature and pressure conditions are selected to achieve the upgrading of the hydrocarbons fed to a product of LPG, a relatively high octane gasoline and/or a diesel or jet fuel with the desired pour point and pour point. Operating conditions can vary considerably depending on the specific hydrocarbon fraction being processed and the requirements for one or more fractions of LPG, high octane gasoline, or diesel or jet fuel.

The tests described below were performed in a microreactor with system components that allowed rapid catalyst preparation and pretreatment, easy adjustment of reaction conditions, feed gas mixtures, and regeneration conditions without disrupting the catalyst. The reactor is 400 mm long and made of 304 stainless steel tube with an outer diameter of 9.525 mm. An annular thermowell with an outer diameter of 3.175 mm made of 304 stainless steel is located along the entire length of the catalytic bed arranged in the tube. The tube contains a 3 ml catalyst bed (length 90 mm) which is placed in the middle of a 300 mm vertical tube furnace. The catalyst bed is held in place with vycor glass wool, while the voids in the reactor supporting the bed are filled with Pyrex tube Rashig rings. The temperature is maintained by a proportional control device with a thermocouple located in the furnace wall near the reactor. During use, hydrogen and carbon monoxide are mixed together, the mixture is compressed and passed through activated carbon columns (this operation is necessary to remove traces of iron carbonyl present in the carbon monoxide bomb). A constant gas flow is maintained by a thermal control device for controlling the mass flow of gas, and the pressure in the reactor is maintained by a backpressure regulator located downstream of the reactor and by liquid scrubbers.

I, Catalyst preparation and pretreatment

Catalysts 1 to VI are prepared by co-precipitating the desired metals from hot aqueous solutions (85 to 90°C) of nitrates by slow addition of hot aqueous sodium carbonate solution (>100% excess carbonate).

The precipitate is filtered and washed until the carbonate in the washings disappears. The washed precipitate is dried in a vacuum oven (70 °C) and calcined in air for at least 6 hours. Droplet analysis of the calcined powders for copper, zinc, chromium and sodium shows that the metals are present in the resulting material in the same proportions as in the original nitrate solutions and that sodium is essentially removed. Details relating to these preparations are given in Table 1 below.

The calcined powders (up to 180 μm) were then physically mixed with equal parts by weight of powdered gamma-alumina as the acidic dehydrating component (the gamma-alumina was previously calcined at 538 °C for at least 6 hours) and the mixture was mechanically pelletized and crushed to a size of 600 to 2,000 μm.

Each of the resulting catalysts is then pretreated in a reactor at 204 °C with a reducing stream (0.1 MPa, hourly gas space velocity 1,500 h^), in which the hydrogen partial pressure is slowly increased by increasing the hydrogen content from 0 to 2 vol. % and then up to 8.5 vol. %. The pretreated catalyst is then cooled to a temperature below 49 °C in an inert gas stream, the feed is switched to synthesis gas, while the pressure, space velocity and temperature are maintained as required.

Another catalyst (catalyst VXI) is prepared as follows: a portion of the carbonate-precipitated methanol conversion component used in the catalyst of Example II is calcined at 260°C. A portion of the resulting powder is mixed with an equal weight of gamma-alumina, the mixture is pelletized and ground, and the 600-2000 µm fraction is used, which has a bulk density of 1.05 g/ml. The catalyst obtained has a Cr/(Cu+Zn) ratio of 0.5.

A ml sample of catalyst VII is placed in a tubular reactor and pretreated in the above-mentioned manner and used in the synthesis gas conversion.

Another catalyst (catalyst VIII) is also prepared in the following manner. A solution is prepared by dissolving copper, zinc and chromium nitrate in the amounts given in Table 1 for catalyst VI in 170 ml of water. A solution of 41.5 ml of ammonium hydroxide (29%) is slowly added to the solution to pH 7.5. The resulting precipitate is washed three times with 1 to 1.5 l of water, dried in a vacuum oven (140 °C) and calcined at 357 °C. A portion of the resulting powder is combined with an equal weight of gamma-alumina, the mixture is pelletized and ground, and a fraction with a particle size of 600 to 2,000 pm is used. The resulting catalyst for the production of dimethyl ether differs fundamentally from catalyst VI in that the metals in the catalytic component for the preparation of methanol were precipitated with ammonia instead of carbonate.

Table 1

Metal components of the catalyst for the synthesis of dimethyl ether

Catalyst for p example I II III IV · In VI Solution A water ml 500 200 200 300 200 200 Cr _{( NO3} ) _3.9H2O g 52.7 48.03 74.1 136.3 74.1 92.6 Cu _{( NO3} ) _2.3H2O g 220.1 29.0 22.4 22.4 34.4 22.9 Zn _{( NO3} ) _2.6H2O g 130.5 35.71 27.5 27.5 13.9 9.2 La _{( NO3} ) _2.6H2O g 71.6x - - - - - Solution B water ml 2,500 550 300 300 550 550 Na ₂ CO ₃ g * 400 111.3 54 81.2 100 100 Precipitation temperature °c 90 1 35 to 90 90 85 to 90 90 95 pH of the filtrate - 9.8 8.0 8.36 9.95 9.4 calcination temperature °c 265 338 274 267 338 260 Cr(Cu+Zn) 0.14 0.5 1.0 1.8 1.0 1.8 Cu/Zn 2.06 1.0 1.0 1.0 3.0 3.1 Catalysts for synthesis of di methyl ether will prepare by mixing the above powders are the same

by mass of calcined gamma-alumina x Analysis of the resulting metal-containing mixture shows that its lanthanum content is 3% molar based on the catalytic components for methanol production

As can be seen from Table 1, the catalyst compositions are different. This allows a comparison of the effect of chromium content at similar Cu/Zn ratios and the effect of different Cu/Zn ratios at similar chromium concentrations.

II, Catalyst aging properties

The aging properties of each catalyst can be quantitatively expressed by the slope of the function ln(-ln(l - Q)) and the operating time in days, where Q is the ratio of the observed total synthesis gas conversion to the conversion predicted for thermodynamic equilibrium. The negative value of this slope is the decay rate for the exponential expression of deactivation. This result is obtained when the expression for the decay rate (1)

K = K _Q . e~ ^bt (1) where

K represents the empirically determined reaction rate,

K _Q is the value of this quantity for a fresh catalyst, b is the decomposition parameter and t is the commissioning time, substitute into the empirical reactor model (2) ln(l - X/X _eq ) = - K/GHSV (2) where

X is the observed total molar conversion of H ₂ + CO,

X is the equilibrium value of X and eq ^J

GHSV stands for gas hourly space velocity during reaction.

By substituting the expression from equation (1) for K in equation (2) and linearizing the equation, we arrive at equation (3) ln(- ln(l “ Χ/Χθ^)) = ΙηίΚθ/GHSV) - bt (3)

Table 2 gives details relating to the experimental conditions using the six catalysts described in Table 1, together with the values of the decomposition constants b obtained.

of the integration of the function expressed by equation (3) by the least squares method. In Fig. the value of the decomposition constant against the chromium content in each of the six catalysts is plotted on the X-axis in the form of the Cr/(Cu+Zn) ratio. . 6 is plotted Chrome touch Table 2 Effect of catalyst composition on its aging Catalyst from example I II III IV in VI Reaction conditions using a batch with a molar ratio of _H2 /CO equal to 1: Pressure (MPa) 10.0 . 10.0 10.0 10.0 10.0 , 10.0 Temperature Pcj 316 317 315 314 316 316 Hourly space velocity of gas Eh ^-1 J 3 800 4,000 4 000 4,000 4 000 4,000 Aging time IdnýJ 40 48 38 20 24 42 Catalyst composition: Cr/(Cu+Zn) 0.14 0.50 1.0 1.0 1.8 1.8 Cu/Zn 2.1 1.0 1.0 3.0 1.0 3.1 Catalyst degradation parameter value (least squares correlation) [days, 0.040 0.013 0.030 0.029 0.032 0.027

and

From the above tests it is evident that the catalyst performance is best at a composition given by a Cr/Cu+Zn ratio of from about 0.1 to about 1.0, with a preferred Cr/(Cu+Zn) ratio of about 0.5.

Deviations from this behavior occur in the case of a low Cu/Zn ratio (0.5) and in a commercially available catalyst for methanol conversion prepared by a method other than co-precipitation. A study of the aging data of catalysts containing 25, 50 and 83% Al ₂ °3 also shows that the best aging behavior is shown by the catalyst with 50% alumina. In the case of a catalyst containing a large amount of chromium, there is a large difference in the activity of the catalyst obtained by precipitation with both ammonia and carbonate. The catalyst prepared by precipitation with ammonia (see below and Fig. 7) shows a lower activity and therefore a lower conversion is achieved. For this reason, carbonate precipitation was continued to be used for the preparation of the catalyst.

III. Catalyst regeneration

It was found early in the development of the invention that at temperatures above about 288°C, the aging of the dimethyl ether synthesis catalyst is mainly due to the aging of the methanol synthesis catalyst components. The detailed aging mechanism of methanol synthesis catalysts is not known, but it is believed that the aging has the following causes:

1) coke storage,

2) catalytic phase changes and

3) changes in oxidation state.

A fourth possibility is strong competitive sorption of carbon monoxide, especially when hydrogen-poor feedstocks are used.

In some pioneering studies of copper-based catalysts for the synthesis of dimethyl ether, the following observations were made: 1) a spent catalyst that had been aged for 20 days and had lost 30% of its activity was analyzed to contain 1.3% coke (0.001% of the feed), 2) calcination of this catalyst in air at 538°C for 16 hours resulted in further deactivation, 3) samples of the spent catalyst showed no capacity for chemisorption of carbon monoxide, 4) treatment of the spent catalyst that had undergone aging with hydrogen did not restore its activity. The state in which the active catalyst components exist is not known, but there is evidence that when the copper is reduced to the metallic state, the catalyst is no longer active, just as a catalyst that has been subjected to too harsh an oxidation treatment is not active. It has also been suggested that a catalyst containing copper in the monovalent valence state Cu ¹ is active (see, for example, Journal of Catalysis 56, 407-429 (1979), where a solution of Cu ¹ zinc oxide is reported as the active component of catalysts for low-pressure methanol synthesis based on combinations of Cu/ZnO, Cu/ZnO/Al ₂ O ₂ and Cu/ZnO/Cr _{2 O 2} ). The quantitative absorption of carbon monoxide in solutions of compounds of monovalent copper is also known, see FA Cotton and G. Wilkinson in Advanced Inorganic Chemistry, second edition, Interscience Publ., London, 1968, p. 898.

It can therefore be assumed that the loss of the catalyst's capacity to chemisorb carbon monoxide can be explained by the reduction of monovalent copper.

In an effort to maintain the activity of the catalysts for the synthesis of dimethyl ether, used in accordance with the invention, over a longer period of operation, the following tests were performed.

Attempts were made to oxidize the Cr/Cu/Zn gamma-Al ₂ O ₃ catalyst by connecting a vessel containing pure or diluted oxygen in an inert gas (helium) to the reactor inlet,

An inert gas was then introduced into the vessel, which mixed with the contents of the vessel by diffusion and carried the gas from the vessel to the spent catalyst. After the oxidation treatment, the vessel and the reactor containing the catalyst were brought to the reaction pressure and temperature using helium. The feed for the feed stream was then switched to synthesis gas. Diffusion mixing in the vessel upstream of the reactor resulted in a gradual shift in composition to predominantly synthesis gas in parallel with the reaction at a set temperature. In this system, the frequency, duration, partial pressure and amount of oxygen used for regeneration were varied. Based on these oxidation studies, a regeneration procedure was developed that can reactivate and substantially stabilize the catalysts for the synthesis of dimethyl ether according to the invention.

Initial studies were conducted using a 1:1:1 Cu:Zn:Cr + 50% gamma-alumina catalyst which was aged for 105 days until it provided 20% conversion (catalyst from Example II).

The minimum effective oxygen partial pressure for processing under reactor conditions (10 MPa, hourly gas space velocity 4,000 h using a 100 1/1 catalyst pulse) was determined. Processing was performed at four oxygen partial pressures (0.01,

0.1, 0.4 and 0.2 MPa) at high total pressure (10 MPa) and temperature ranging from 307 to 338 °C. The results achieved are summarized in the first four examples of Table 3 below.

After regeneration, the initial conversion increased to 140%. However, the regenerated catalysts gradually aged, with the conversions achieved gradually decreasing to the level before regeneration within about 10 days. It is therefore clear that at high pressure and a pulse volume of 100 l gas/l catalyst, the oxygen partial pressure is preferably higher than 0.01 MPa, preferably 0.1 to 0.4 MPa or higher.

The total volume of oxygen brought into contact with the catalyst during regeneration was then increased, as shown in the last three examples in Table 3. In these regenerations, the partial pressure of oxygen was maintained at about 0.1 MPa, but the temperature of the oxidation treatment was increased. As is evident from Table 3, the activity of the catalyst decreased, as evidenced by the changes in conversion after the catalyst was treated with oxygen. For example, after 108.8 days of operation, the catalyst was regenerated with oxygen and its activity was regenerated to a conversion level of 47.6%. Then the activity decreased to a conversion level of 37.4 and after 114.0 days of operation, the activity was regenerated by oxidation to a conversion level of 47.1%. After treatment with large volumes of oxygen, as described in the last three examples, the catalyst activity was regenerated to a conversion level of 55.8,

56.6 and 44.8%.

Table 3

Mass balance before and 16 hours after treatment of the catalyst with oxygen

Oxygen volume

[N 1/1 catalyst] 10.0 100 400 200 2 100 5 000 500 Oxygen partial pressure [MPa] 0.01 0.1 0.4 0.2 0.1 0.1 0.1 Total pressure (Helium diluent) QlPa] Hourly space velocity 10.0 10.0 10.0 10.0 0.1 0.1 0.1 oxygen gas Ch ^ ] 4 000 4 000 4 000 4 000 100 900 900 Temperature L°cl 307 338 309 321 293 458 514 Reaction conditions: Operation time [days] 104.0 104.8 105.8 108.8 112.8 114.0 119.0 Temperature in the reactor Γ°^] 315 315 317 317 316 318 317 Reactor pressure QíPa] Hourly space 10,003 9,935 9,935 9,935 9,935 9,935 9,935 gas velocity Ch ] 4 000 4 000 4 000 4 000 4 000 4 000 4 000 Weight hourly space- level speed Ch 3.18 3.18 3.18 3.18 3.18 3.18 3.18 kg/l/h 2.68 2.68 2.68 2.68 2.68 2.68 2.68 Conversion [% mol] : ^H 2 20.8 25.8 26.6 49.2 35.9 51.1 31.4 WHAT 18.2 20.8 19.6 46.0 38.9 43.1 23.5 Total 19.5 23.3 23.1 47.6 37.4 47.1 27.5 Reaction conditions: Operation time [days] 120.0 122.0 123.0 125.0 126.0 127.0 128.0 Temperature in the reactor C°c3 371 316 317 316 315 316 315 Reactor pressure [MPa] Hourly space 9,935 9,935 9,935 9,935 9,935 9,935 9,935 gas velocity C ^h ^] ⁴ 000 4 000 4 000 4 000 4 000 4 000 4 000

Table 3 continued

Hourly rate of weight*

rhythmic speed [h 3.18 3.18 3.18 3.18 3.18 3.18 3.18 kg/l/h 2.68 2.68 2.68 2.68 2.68 2.68 2.68 Conversion [% mol]: ^H 2 46.8 32.6 60.8 48.0 59.6 44.4 47.3 WHAT 38.6 24.7 50.8 43.2 53.5 42.0 42.2 Total 42.7 28.6 55.8 45.6 56.6 43.2 44.8

The second regeneration study was performed for a catalyst with a metal ratio of Cu:Zn:Cr 1:1:1, containing an additional 25% alumina. The catalyst was allowed to age in service to a conversion level of 29%. During 11 oxidation pulse cycles (0.1 MPa 0^ at 10 MPa and a volume of 100 l/l of catalyst every fifth day), the activity was maintained at a conversion level above 45%. Between days 95 and 101 and days 104 and 108, daily oxidation treatment was able to maintain a minimum mean conversion of 44% (+ 2.6%). Between days 101 and 104, the volume of the oxidation pulse was reduced to 25 l/l of catalyst. During this three-day period, the conversion dropped from 42 to 36%, clearly indicating that oxygen treatment under harsher conditions (100 l/l catalyst at 0.1 MPa O ₂ ) is more advantageous.

After 68 days of operation, a hydrogen treatment was performed after the oxidation pulse to activate or pretreat the fresh catalyst. Conversion data after this treatment showed that there was little, if any, regeneration effect. The hydrogen treatment used was apparently sufficient to deactivate those sites that had been regenerated by the oxidation step.

During the last two cycles (107 and 108xden) the oxidation treatment was carried out at a total pressure of 0.1 MPa with pure oxygen instead of at 10 MPa with 1% oxygen. No major differences in the activity of the regenerated catalyst were found. If the initial partial pressure of oxygen in the pulse was maintained at about 0.1 MPa, it was shown that any practical total pressure (0.1 to 10 MPa) could be maintained with inert gases.

In the third study, a fresh catalyst (Cu:Zn:Cr = 1:1:1 + 50% alumina) was first subjected to three-day and daily regeneration cycles (0.1 MPa _O2 , 10 MPa total pressure, volume

100 l 0^/100 l catalyst, temperature 293 to 343 °C, hourly oxygen space velocity up to 600 h ) under reaction conditions involving high space velocity (hourly gas space velocity 4,000 h , temperature 316 °C, 10 MPa, H ₂ /CO - 1). Then, after a period without regeneration, a daily regeneration regime was applied, under reaction conditions involving low space velocity (hourly gas space velocity 2,000 h ~' ^Í , temperature 316 °C, total pressure 10 MPa, H ₂ /CO 1 ratio). The period of daily regenerations was long (24 cycles) and was carried out under the conditions: total pressure 0.1 MPa, _O2 pressure 0.1 MPa, catalyst volume 100 l _O2 /l, temperature 293 to 343 °C, oxygen hourly space velocity 70 to 600 h\ This shows that stable operation at essentially equilibrium conversion can be maintained by daily regenerations. The results of this test are shown in Fig. 8.

In another test, catalyst VII was kept in operation for 50 days with periodic oxidation regenerations. The nominal reaction conditions were: pressure 10 MPa, temperature 316 °C, molar ratio

H„/CO in synthesis gas 1:1 and hourly gas space velocity 4,300 h for the first 28 days, — 1 — 1

000 h on the 29th day and 2,000 h from 30 to 49.5 days of operation. Periodic oxidation regeneration of this catalyst was carried out by the following two procedures:

A. For the first 28 days of operation, oxidation regeneration was carried out according to this regulation:

1. The reactor is purged with helium at a pressure of 10 MPa.

2. Elute at a pressure of 10 MPa with a 1% mixture of oxygen in helium placed in a reservoir, at a volume of

1 per liter of catalyst (i.e. 100 1 O, at standard pressure per liter of catalyst), an hourly gas space velocity of 7,000 h (the catalyst temperature was maintained between 271 and about 316 °C during this stage).

3. Flushed with helium at a pressure of 10 MPa.

4. Synthesis gas is fed into the reactor to restart the reaction at 316°C and 10 MPa, proceeding by feeding 100% synthesis gas into the reactor via a helium buffer, the feed rate being 100 1/1 catalyst at a gas hourly space velocity of 4,306 h\

B. Oxidative, regeneration after 28 days was carried out according to this regulation.

1. The reactor is depressurized to atmospheric pressure and purged with helium.

2. Elute with a pulse of 100% oxygen from the reservoir, with a pulse volume of 100 liters/liter of catalyst (i.e. 100 liters of oxygen at standard pressure per liter of catalyst), a pressure of 0.1 MPa is set with helium, a gas hourly space velocity of 600 h and a temperature of 260 to 316 °C.

3. The system is flushed with helium and the pressure in the system is adjusted to 146 MPa with helium.

4. The reactor is again fed with 100% synthesis gas, as in stage 4 of the regime described in regulation A, and then the hourly gas space velocity is set to the desired value of 2,000 (or 3,000) h \

It was found that while three-day cycles were not effective in maintaining steady-state operation at an hourly gas space velocity of 4,000 h ^x , using daily low-sharpness regeneration cycles, stable operation was maintained for 1 week with 56% conversion (70% of the equilibrium value). Statistical correlation of these data indicated that this was essentially a “steady-state” activity.

At a lower dosing rate (gas space hourly velocity of 2,000), aging occurred without regeneration. However, daily regeneration allowed maintaining essentially equilibrium conversion (79% versus 80%) for 20 days of operation (24 cycles). A mass balance was performed at the end of each cycle. Aging was not statistically significant.

The main conclusion of this catalyst study is that values above 50% of the equilibrium conversion value and values substantially up to the 80% conversion level can be maintained with stable behavior of the dimethyl ether synthesis catalyst if periodic or relatively continuous catalyst regeneration is performed under low severity conditions.

In the last test, catalyst VIII (precipitated with ammonia) was kept in operation for several days under the following nominal conditions: pressure 10 MPa, temperature 316 °C, molar ratio H2/CO in the reported synthesis gas 1:1 and hourly space velocity of the gas phase 4,000 h. The activity of catalyst VIII compared to the activity of catalyst VI is shown in Fig. 7.

IV, Effect of catalyst preparation on regeneration

In order to further determine the extent to which the presence of chromium in the catalytic component for methanol synthesis of the catalyst for dimethyl ether synthesis is of decisive importance and to determine the influence of the co-precipitation technology, the response of different catalytic systems to oxidative regeneration under mild conditions was studied. Details related to the composition of the investigated catalysts are given in Table 4. The catalysts for dimethyl ether synthesis were obtained by pelletizing physical mixtures of equal parts by weight of gamma-alumina and the component for methanol synthesis. Catalysts IX and X were prepared by co-precipitation of Cu/Zn/Cr and Cu/Zn/Al mixtures with an excess of sodium carbonate (Na2CO^) solution to a final pH of 9. The metals were used for precipitation in the form of nitrate solutions. Calcination after precipitation was carried out at 260 °C. Catalyst XI was prepared by carbonate co-precipitation with a binary mixture of Cu/Zn and calcination at 343 °C. Catalyst XII was prepared by carbonate co-precipitation of a binary Cu/Zn mixture, calcination at 343°C, and mixing in aqueous suspension with Al ₂ O ₃ under high shear in a Wang mixer according to the procedure described in U.S. Patent No. 3,790,505. Catalyst XIII was prepared by carbonate precipitation of binary Cu/Cr and Zn/Cr systems, and subsequent mixing of the powders formed from these precipitates by calcination at 260°C. Other details of the preparation of these catalysts were the same as in the preparations described above. For example, all catalysts were activated with dilute hydrogen at low temperature to avoid undesirable exotherms.

Synthesis gas conversion on these catalysts was monitored using the apparatus described above (316 °C, 10 MPa, H ₂ /CO = 1 and a gas space hourly velocity of 4,000 h ^{-1 )} with the exception of catalyst IX, where a gas space hourly velocity of 4,000 h ^-1 was used up to day 28 and then 2,000 h %. Periodic oxidation regeneration was carried out by procedure B described above with the difference that the temperature was 288 to 343 °C.

The behavior of the basic Cu/Zn/Cr system during regeneration is illustrated in Fig. 10. Like the data in Fig. 8, these data also show that the activity of catalyst IX can be regenerated by oxidative regeneration. Constant conversion can be achieved with sufficiently frequent periodic regenerations.

The effect of simultaneous precipitation of modifying metals (Cr and Al) evident from Fig. 10 and 11 can be compared with the situation when the modifier was absent (Fig. 12). The initial activity of the catalysts (expressed in conversion) is quite similar. However, when the modifiers were not present, even the first regeneration of catalyst XI did not bring any benefit. It is therefore clear that the Cu/Zn-based components must be stabilized by structural promoters.

Table4

Components of the catalyst for the synthesis of dimethyl ether, whose task is to synthesize methanol

Catalyst Composition of the component for methanol synthesis and its preparation

IX Cu/Zn/Cr = 1/1/1, carbonate precipitation

X Cu/Zn/Al = 1/1/1, carbonate precipitation

XI Cu/Zn = 1/1, carbonate precipitation

XII Cu/Zn/Al - 1/1/1.5 Cu/Zn carbonate precipitation,

A1 ₂ O ₃ interrupted by high shear mixing

XIII Cu/Cr + Zn/Cr = 1/0.5 + 1/0.5

Note: base case: chromium stabilizer ⁴ replaced by aluminum; aluminum and chromium stabilizer omitted; mixing with alumina performed under high shear conditions; influence of separate co-precipitation of active metals with chromium

The introduction of the modifying aluminum by physical mixing is illustrated in Fig. 13. The activity and aging of the catalyst (catalyst XII) were close to those of the unmodified catalyst XI (Fig. 12). The catalyst that succumbed to aging could not be regenerated. It is clear that the modifier must have been co-precipitated with the active metals.

The case where half of the modifying chromium is co-precipitated with copper and the other half with zinc, after which the resulting calcined powders are combined with a dehydrating component to form a catalyst for dimethyl ether formation (catalyst XIII) supports this conclusion.

Although this catalyst was less active, it was regenerable, as illustrated in Fig. 14 . The significantly reduced activity of this combination is consistent with the hypothesis that the active sites for methanol synthesis form a solid solution of copper in the zinc oxide phase.

The coprecipitated Cu/Zn/Al catalysts for dimethyl ether synthesis initially have a higher activity for regeneration than similar chromium catalysts (compare Fig. 11 and Fig. 10). However, after 40 days of operation and 4 oxidation treatments, the aluminum-modified catalysts are no longer regenerable.

The coprecipitated Cu/Zn/Cr catalysts for the synthesis of dimethyl ether according to the invention exhibit remarkable stability to regeneration. A steady high conversion was observed at a gas hourly space velocity of 2000 h ^-1 for 80 days and 45 cycles. Even at a gas hourly space velocity of 4000 h ^-1, the catalyst was oxidatively regenerable after 125 days and 5 cycles.

Fig. 9 illustrates the data obtained to evaluate the range of severity of the conditions (temperature) during oxidative regeneration. The conversion was carried out at a temperature of 316 °C, a pressure of 10 MPa and an hourly gas space velocity of 4,000 h ¹ using a synthesis gas with a molar ratio of H^/CO = = 1 and a catalyst based on Cu/Zn/Cr = 1:1:1. The following regeneration conditions were used:

Experiment Temperature (°C) Pressure (MPa) O ₂ molarJ

and x 316 10.0 0.1 b 316 10.0 1 c 316 10.0 4 d 316 10.0 2 e xx 293 0.1 100 f 399 0.1 100 g 454 0.1 100 h 510 0.1 100 and 316 0.1 100 xxx j 293 0.1 100

x catalyst aged 104 days until conversion dropped below 20% xx in runs a to e (122 days) gas hourly space velocity 4,300 h xxx followed by hydrogen treatment at 204 °C.

As can be seen, the regeneration reaction temperature varied from 293 to 510 °C. Regeneration at temperatures above 510 °C resulted in a likely irreversibly reduced catalyst activity.

V. Effect of simultaneous dosing of water together with hydrogen-poor synthesis gas.

Synthesis gas poor in hydrogen (molar ratio H ₂ /CO < 1), which is obtained in current and sometimes preferred more economical gasifiers, can be used with considerable advantage in the selective conversion to dimethyl ether using suitable synthesis catalysts. In such an operation, water can be brought into contact with the selected dimethyl ether synthesis catalyst simultaneously with the synthesis gas with a low H ₂ /CO ratio. On the catalyst, due to the activity of its components for catalyzing the conversion of synthesis gas to water gas, water with the same amount of carbon monoxide is converted to hydrogen plus carbon dioxide. As illustrated by the operation under equilibrium conditions in Figs. 1 and 2, the utilization of the synthesis gas and/or the molar ratio of H ₂ /CO of the gas leaving the dimethyl ether synthesis stage can be improved in this way.

The gas with an improved _H2 /CO ratio obtained from the dimethyl ether synthesis stage due to the simultaneous addition of water necessarily contains other gases, mainly carbon dioxide. In view of the significant imbalance of the conversion reaction leading to dimethyl ether by carbon dioxide as a diluent, recycling this gas containing carbon dioxide to the dimethyl ether synthesis stage, accompanied by the necessary purification of the recycled gas, may not be the best way of utilization.

The regenerated unreacted synthesis gas stream containing synthesis gas with a correspondingly improved _H2 /CO ratio can alternatively be introduced into conversion systems in which the conversion is not limited by an equilibrium state (as is the case in dimethyl ether synthesis) and which can (perhaps with lower selectivity) utilize the unreacted synthesis gas stream. A secondary reaction system that is particularly suitable for the conversion of this synthesis gas containing carbon dioxide is a Fischer-Tropsch conversion unit using an iron, cobalt or ruthenium-based synthesis catalyst. Such a secondary catalytic synthesis system allows for optimal utilization of the synthesis gas recovered from dimethyl ether synthesis with minimal or no intermediate removal of carbon dioxide. Such a system is illustrated in Fig. 5.

The synthesis gas stream with an improved _H2 /CO ratio from the dimethyl ether synthesis unit is particularly suitable for further processing by the Fischer-Tropsch process using an iron-based catalyst. A publication published in Industrial and Engineering Chemistry - Process Research and Development, 15 (4) 1976 states that it is useful to increase both the _H2 /CO ratio ( _H2 /CO <6) and the carbon dioxide concentration at the inlet to the Fischer-Tropsch reactor. By reducing the ratio (Ρ^/Ρ^ ) ³ , i.e. by increasing the _H2 /CO ratio, carbon formation is reduced. As Pqq increases, the selectivity of the conversion to hydrocarbons towards methane also decreases. The inclusion of this

The Fischer-Tropsch stage, as a second stage after the dimethyl ether synthesis stage, in which the H ₂ /CO ratio and the C0 ₂ content in the gas regenerated in the first stage are increased due to the simultaneously fed water, is therefore beneficial.

The _H2 /CO ratio in the synthesis gas recovered or discharged from the dimethyl ether unit as unreacted synthesis gas may not be high enough to allow the synthesis gas to be processed in a secondary unit when operating without co-fed water. Synthesis gas upgrading can be carried out in a separate reaction unit to convert the synthesis gas to water gas, but is more economically accomplished by co-feeding water with the hydrogen-poor synthesis gas to the dimethyl ether unit. Secondary systems including a Fischer-Tropsch synthesis unit operating with iron, cobalt or ruthenium catalysts typically require synthesis gas with a molar ratio of H2/CO of 1 or higher, which is obtained as the effluent from the dimethyl ether synthesis stage. The amount of water to be fed into the dimethyl ether synthesis together with the synthesis gas obtained in the gasification step is dictated by the actual synthesis reaction leading to dimethyl ether, as described below. The amount of water fed together with the synthesis gas should also be limited by an upper limit.

In the case of a feed consisting of synthesis gas with a molar ratio (r) H ₂ /CO lower than 1, water can be dosed into the dimethyl ether synthesis reactor in the amount given by equation (4)

XX ¹ ' ⁴ >

where

S represents the molar ratio of H ₂ O/CO and r' represents a stoichiometric parameter that has an arbitrary value limited by the relation r 4r”^l

The molar ratio of _H2 /CO in the gas discharged from the reactor will increase with increasing S value.

The limiting values for S can be expressed under conditions of ideal behavior as follows:

1) The upper value of the ratio S, S</l is given by the fact that in the case of S = 1, carbon monoxide would be completely converted by water into hydrogen and carbon dioxide, so that no carbon monoxide would remain available for the subsequent synthesis of ether.

2) The stoichiometric parameter of the synthesis gas r' is of great importance. The stoichiometry of the reaction to dimethyl ether assumes a molar ratio of H ₂ /CO in the feed r' = 1.0. When the molar ratio of H ₂ /CO in the feed r is less than 1, water fed simultaneously in a molar amount of o - * - 1 - r ^b 1 + r' 2 would provide an effective ratio of H ₂ /CO equal to 1 if all the water fed simultaneously were completely converted with the available carbon monoxide to hydrogen and carbon dioxide.

3) The requirement for a lower bound on the value of r' r'(r is an expression of the requirement that the numerator of the left-hand side of relation (4) r~- r must be positive.

+ r'

The limitations and discussion above are of course idealized. The actual values depend on the complex equilibrium and dynamic behavior in the reactor, and the above considerations are complicated by catalyst bypass and mixing downstream of the reactor. For a specific reactor configuration and synthesis parameters, extensive variables such as the synthesis gas conversion or the _H2 /CO ratio at the reactor outlet can be correlated with the water feed parameter (S). Using these correlations, a desired value for S can be determined within the range given by equation 4.

Such a correlation is illustrated by the following series of tests.

Two additional catalysts XIV and XV ^were prepared in two parts. First, the methanol synthesizing catalyst component was prepared by co-precipitating the three metal components from a stirred aqueous solution of the nitrates of these metals by adding an excess of hot (85-90°C) sodium carbonate solution. The resulting precipitate was washed, dried, and calcined at about 269°C.

These methanol synthesis components were then ground into powders and each was mixed with an equal weight of powdered gamma-alumina.

Catalyst XIV contained a methanol synthesis component made from equal amounts of copper, zinc and chromium. Catalyst XV contained a methanol synthesis component made from equal amounts of copper, zinc and aluminium.

Both catalysts were first activated at 0.1 MPa and 204 °C by contacting them with a gas stream consisting of hydrogen and an inert gas, in which the hydrogen content was slowly increased to 2% vol. and then to 8.5% vol. The catalysts were then operated for a long time under synthesis operating conditions with periodic oxidative regeneration. The oxidative regenerations were carried out by flushing the reactor with an inert carrier gas and then passing a pulse of oxygen at a volume of 100 l (standard pressure) per liter of catalyst through the catalyst at atmospheric pressure. The temperature during this treatment was 288 to 343 °C and the space velocity was sufficiently low to avoid thermal damage to the catalyst. After the oxygen pulse, the reactor was flushed again with an inert gas, which was gradually replaced by synthesis gas while simultaneously ramping up to reaction conditions.

Catalyst XIV operated with daily regenerations under the following nominal conditions:

316 °C, 5 to 5.4 MPa, hourly gas space velocity (relative to synthesis gas) 1,000 h ¹ , molar ratio H ₂ /CO 2:3, various amounts of water introduced. As shown in Fig. 3, the amount of water introduced together changed during the experiment. In Fig. 4, the results of this experiment are correlated with the amount of water introduced together with synthesis gas, relative to the unit amount of carbon monoxide (S) introduced. Two variables were used for the correlation, namely the molar ratio H ₂ /CO at the reactor outlet (Fig. 4A) and the total conversion of synthesis gas, i.e. reacted H ₂ +CO (Fig. 4B). The measured values are compared with the values predicted for conversions under thermodynamic equilibrium conditions under the respective dosing conditions. Results such as those shown in Fig. 4 can be obtained for any particular dimethyl ether synthesis system.

Catalyst XV was contacted for a long time at 316 °C with synthesis gas with a molar ratio of H ₂ /CO 2:3 at a water dosing level together with synthesis gas of 0.25 mol HjO per 1 mol CO at a pressure of 5.4 MPa and with periodic oxidative regeneration of the catalyst. The behavior of the catalyst 18 hours after regeneration and after 30 days of contact is described in Table 5. The molar ratio of H^/CO in the gas leaving the reactor of 0.85 shows that the gas has been enriched with hydrogen. A comparison of the observed ratio of H ₂ /CO in the gas leaving the reactor and the conversion in this experiment using catalyst XV shows a deviation from equilibrium that is similar to that found by correlation of the results in the case of catalyst XIV.

Table 5

Dimethyl ether synthesis products 18 hours after regeneration using catalyst XV

Handle:

Molar ratio H ₂ /CO 0.67

Molar ratio H ₂ O/CO 0.25

Reaction conditions:

Operating time |days 30.0

Reactor temperature °C I 316

Reactor pressure (MPaj 5.4

Hourly gas space velocity fh 1,200

Hourly mass, space velocity Γ h 0.9 8

Mass-volume space velocity Ikg.l \h ^7 0.95

Product composition from the reactor mass.

^H 2

CO dimethyl ether ch ₃ oh ch ₂ ch ₂₊

Total conversion

Selectivity: D/(D+M) l%J Yield related to fkg.l' ¹ .h ^-11

1.9

31.3 45.5

0.56

19.3 1.3 0.10 0.04

58.3

0.195 on space and time

In the last experiment, the Fischer-Tropsch process was used in a tubular reactor using a catalyst consisting of compressed powdered products of Fe ₃ O ₄ and gamma-alumina in a weight ratio of 1:10. The reaction conditions were as follows: molar ratio of H ₂ /CO in the batch 1:1, pressure (H ₂ +CO) = 2.7 MPa, mass hourly space velocity (relative to H ₂ +CO and metal) 10.3 h ^_1 , temperature 320 °C. The following results were achieved: conversion of H ₂ + CO = 89%, composition of hydrocarbon products: methane 21%, C ₂ to C ₄ hydrocarbons 37% and C ₅₊ hydrocarbons 42%. Similar results could be expected if the reactor feed contained synthesis gas with a molar ratio of H ₂ /CO = 1 and carbon dioxide as a diluent, provided that the partial pressure (H ₂ +CO) was maintained at a similar level of 2.7 MPa.

The regenerated unreacted synthesis gas leaving the dimethyl ether production reaction system operating with the catalyst XIV described above can be adjusted to have a _H2 /CO molar ratio of 1 or higher by simultaneously introducing an appropriate amount of water into the dimethyl ether production reactor. It can be seen from the graph in Fig. 4A that the amount of water introduced onto the catalyst that would result in a regenerated gas having a H2/CO molar ratio of 1 would correspond to a H2O/CO molar ratio of 0.31. In the specific case of the two-stage operation according to the present invention illustrated in Fig. 5, the catalyst XIV would operate in the above manner using a feed water ratio of 0.31. This would yield a dimethyl ether product at about 47% synthesis gas conversion and the regenerated synthesis gas stream would have a molar ratio of H2/CO of about 1, which is a ratio suitable for feeds useful in the Fischer-Tropsch process.

Claims (14)

SUBJECT OF THE INVENTION Process for the production of dimethyl ether, optionally together with methanol, by synthesis gas conversion, characterized in that the synthesis gas having a molar ratio of H 2 / CO in the range of 0.5 to 3.0 is introduced, optionally in the presence of water, at 232 up to 399 ° C in contact with a catalyst comprising co-precipitated Cr, Cu and Zn components and an acidic dehydrating component selected from gamma-alumina, silica-alumina, crystalline zeolite type ZSM-5 with high silica content, phosphate, titanium dioxide in in combination with silica, rare earth and clay, in which the atomic ratio Cr / (Cu + Zn) is 0.1 to 1.0 and the acid dehydration component content is 5 to 95% by weight, after which the catalyst is regenerated by contact with a gas containing oxygen at 38-538 ° C. 2. The process of claim 1 wherein the catalyst is recovered at a temperature of from 288 to 454 ° C. 3. The process of claim 1 wherein the catalyst is regenerated at a temperature of from 293 to 343 ° C. 4. A process according to any one of claims 1 to 3, characterized in that after regeneration the catalyst is contacted with the synthesis gas at a temperature of 177 to 232 [deg.] C. and then gradually increased up to the temperature of the conversion process. 5. A process according to any one of claims 1 to 3, wherein after regeneration the catalyst is contacted with the synthesis gas diluted with inert gas at a temperature of 177 to 232 [deg.] C. and then the temperature and concentration of the synthesis gas are raised to the operating values. in the conversion process. 6. The process of claims 1 to 5 wherein the atomic Cr / (Cu + Zn) ratio in the catalyst is 0.25 to 0.75 and the atomic Cu / Zn ratio is u to 5 to 3.0. 7. A process according to claim 1, wherein the acid dehydration component of the catalyst is gamma-alumina. 8. The process of claim 7 wherein the gamma-alumina is 50% by weight of the catalyst. ~ 9. Process according to claims 1-8, characterized in that the synthesis gas feedstock having a molar ratio H ₂ / CO ranging from 0.5 to 2.0. 10. The process of claims 1 to 8, wherein the synthesis gas feed has a molar ratio of < RTI ID = 0.0 &gt; I / CO &lt; / RTI &gt; 11. The process of claim 10 wherein water is added to the catalyst along with the synthesis gas. 12. The process of claim 11 wherein the amount of water is up to 50 moles per 100 moles of carbon monoxide in the synthesis gas. * 13. The method of claim 10 to 12 for preparing a product comprising dimethyl ether and methanol, characterized in that it applies to the conversion of synthesis gas with a molar ratio H ₂ / CO (r) smaller than 1, and water at a defined relationship R'- R / S <1 1 + r ~ where S represents the molar ratio H ₂ O / CO and ar 'represents a stoichiometric parameter having an arbitrary value limited by the relation r ^z r 1. 14. The method of claim 13 wherein r ' is 1.0.