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Title:
PROCESS FOR THE DIRECT CONVERSION OF OXYGENATED COMPOUNDS TO LIQUID HYDROCARBONS HAVING A REDUCED AROMATIC CONTENT
Document Type and Number:
WIPO Patent Application WO/2010/097175
Kind Code:
A1
Abstract:
The present invention relates to a double stage process for the direct conversion of oxygenated compounds to liquid hydrocarbons with a low aromatic content, carried out in two tubular catalytic fixed bed reactors in series wherein said first reactor operates at a higher temperature than the operating temperature of said second reactor, said process comprising the following steps: feeding said oxygenated compounds comprising alcohols and/or ethers represented by the formula (CnH2n+1) -O- (CmH2m+1) wherein n = 1-4 and m = 0-4 to a first reactor, converting said oxygenated compounds, with a first catalyst, to a mixture of hydrocarbons comprising light olefins, typically rich in olefins Cn with n = 2-4, and water; feeding said mixture of water and hydrocarbons comprising light olefins, coming from said first reactor to the second reactor, converting it in said second reactor, with a second catalyst, to a mixture of liquid hydrocarbons which includes heavy olefins Cn with n ≥ 5, non-condensable hydrocarbons and water; separating said hydrocarbon mixture into two streams: one stream containing a mixture of liquid hydrocarbons and water, which separate into an organic phase and into an aqueous phase, the last eventually recycled to said first and/or second reactor, wherein said liquid hydrocarbons comprise heavy olefins Cn with n ≥ 5 and dissolved non-condensable compounds; one stream which includes a mixture of non-condensable hydrocarbons which are recycled to feed said first reactor.

Inventors:
RAMELLO, Stefano (Via Alfieri, 15/D, Novara, I-28100, IT)
PAPARATTO, Giuseppe (Via Vasari 7, Cinisello Balsamo-Milan, I-20092, IT)
RIVETTI, Franco (Via Oglio 28, Milano, I-20139, IT)
Application Number:
EP2010/000962
Publication Date:
September 02, 2010
Filing Date:
February 10, 2010
Export Citation:
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Assignee:
ENI S.P.A. (Piazzale E. Mattei 1, Roma, I-00144, IT)
RAMELLO, Stefano (Via Alfieri, 15/D, Novara, I-28100, IT)
PAPARATTO, Giuseppe (Via Vasari 7, Cinisello Balsamo-Milan, I-20092, IT)
RIVETTI, Franco (Via Oglio 28, Milano, I-20139, IT)
International Classes:
C10G3/00; C07C1/20; C07C2/00; C10G50/00
Domestic Patent References:
2006-07-27
2007-04-19
Foreign References:
US4025576A1977-05-24
US4547602A1985-10-15
US5672800A1997-09-30
US3894103A1975-07-08
US3894104A1975-07-08
US3894106A1975-07-08
US3894107A1975-07-08
US4035430A1977-07-12
US4058576A1977-11-15
US4304951A1981-12-08
US4025576A1977-05-24
US4476338A1984-10-09
US4482772A1984-11-13
US4497968A1985-02-05
US4506106A1985-03-19
US4543435A1985-09-24
US4547602A1985-10-15
US4579999A1986-04-01
US4689205A1987-08-25
US4898727A1990-02-06
US4899002A1990-02-06
US4929780A1990-05-29
US5045287A1991-09-03
US5177279A1993-01-05
US4992611A1991-02-12
US6372949B12002-04-16
US3702886A1972-11-14
EP0172686A11986-02-26
Other References:
A. CORMA ET AL. CHEM. REV. vol. 106, 2006, pages 4044 - 4098
KOKOTAILO G.T.; LAWTON S.L.; OLSON D.H.; MEIER W.M.: 'Structure of synthetic zeolite ZSM-5' NATURE vol. 272, 1978, pages 437 - 438
OLSON D.H.; KOKOTAILO G.T.; LAWTON S.L.; MEIER W.M.: 'Crystal Structure and Structure-Related Properties of ZSM-5' J. PHYS. CHEM. vol. 85, 1981, pages 2238 - 2243
CH. BAERLOCHER; W.M.MEIER; D.H. OLSON: 'Atlas of Zeolite Framework Types', 2001, ELSEVIER
Attorney, Agent or Firm:
DE GREGORI, Antonella et al. (Barzano' & Zanardo Milano S.p.A, Via Borgonuovo 10, Milan, I-20121, IT)
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Claims:
CLAIMS

1. A double stage process of direct conversion of oxygenated compounds to liquid hydrocarbons with low aroraatics content, being carried out in two tubular catalytic fixed bed reactors in series wherein said first reactor works at a higher temperature than the operating temperature of said second reactor, said process comprising the following steps:

feeding to a first tubular catalytic fixed bed reactor said oxygenated compounds comprising alcohols and/or ethers represented by the formula

(CnH2n+i) -O- (CmH2m+i) where n=l-4 and m=0-4 to the first reactor, converting said oxygenated compounds, with a first catalyst, to a mixture of hydrocarbons comprising light olefins, typically rich in olefins Cn with n=2-4, and water;

feeding to a second tubular catalytic fixed bed reactor said mixture of water and hydrocarbons comprising light olefins coming from said first reactor to second reactor, converting it into said second reactor, with a second catalyst, to a mixture of liquid hydrocarbons which include heavy olefins Cn with n ≥ 5, non condensable hydrocarbons and water; ■ separating said hydrocarbon mixture into two streams : o a stream which includes a mixture of liquid hydrocarbons and water, which separates into an organic phase and into an aqueous phase, the last eventually recycled to said first and/or second reactor, wherein said liquid hydrocarbons comprise heavy olefins Cn with n ≥ 5 and dissolved non condensable compounds , ; o a stream which includes a mixture of non condensable hydrocarbons which are recycled to feed said first reactor.

2. The process according to claim 1 wherein the reaction temperature of said first tubular catalytic reactor is in the range of from 3500C to 6000C, wherein the oxygenated compound is fed to said first tubular catalytic reactor at a space velocity

(WHSV) , expressed as weight of fed mass on mass unity of catalyst, in the range from 0.1 h-1 to 100 h-1, wherein the pressure in said first tubular catalytic reactor is kept in the range of from 1 bar absolute to 10 bar absolute.

3. The process according to claim 1 wherein the reaction temperature of said second tubular catalytic reactor is in the range of from 2500C to 4000C, wherein said mixture comprising light olefins is fed to said second tubular catalytic reactor at a space velocity (WHSV) , expressed as weight of fed mass on mass unity of catalyst, is in the range of from 0.1 h-1 to 100 h-1, wherein the pressure in said second tubular catalytic reactor is in the range of 1 bar absolute to 10 bar absolute.

4. The process according to claim 2, wherein the reaction temperature into said first tubular catalytic reactor is in the range of from 4000C to 5500C, wherein the oxygenated compound is fed to said first tubular catalytic reactor at a space velocity (WHSV) , expressed as weight of fed mass on mass unity of catalyst, is in the range of from 0.5 h-1 to 10 h-1, wherein the pressure in said first tubular catalytic reactor is in the range of from 1 bar absolute to 5 bar absolute. 5. The process according to claim 4, wherein the oxygenated compound is fed to said first tubular catalytic reactor at a space velocity (WHSV) , expressed as weight of fed mass on mass unity of catalyst, is in the range of from 1 h-1 to 5 h-1. 6. The process according to claim 3, wherein the reaction temperature into said second tubular catalytic reactor is in the range of from 3000C to 3500C, wherein said mixture comprising light olefins is fed to said second tubular catalytic reactor at a space velocity (WHSV) , expressed as weight of fed mass on mass unity of catalyst, is in the range of from 0.5 h-1 to 10 h-1, wherein the pressure in said second tubular catalytic reactor is in the range of from 1 bar absolute to 5 bar absolute . 7. The process according to claim 6, wherein said mixture comprising light olefins is fed to said second tubular catalytic reactor at a space velocity

(WHSV) , expressed as weight of fed mass on mass unity of catalyst, is in the range of from 1 h-1 to 5 h-1. 8. The process according to claims from 1 to 7, wherein the mixture of non condensable hydrocarbons is eventually recycled to said second reactor. 9. The process according to claims from 1 to 8 , wherein the reaction product is further separated in order to recover the dissolved non condensable hydrocarbons, which are eventually recycled to feed said first and/or second reactor. 10. The process according to one of the previous claims wherein said catalysts are calcined. 11. The process according to claims from 1 to 9 wherein said oxygenated compound is diluted with inert gases and/or water. 12. The process according to claim 11, wherein the molar ratio between the inert gas and/or the water and the oxygenated compound is in the range of from

0.01 to 10.

13. The process according to claim 12, wherein the molar ratio between the inert gas and/or the water and the oxygenated compound is in the range of from 0.05 to 5.

14. The process according to claims from 1 to 9 and 11 wherein said light olefins Cn with n=2-4 and mixtures thereof are co-fed to said oxygenated compound and the ratio by weight between the light olefin and the oxygenated compound is in the range of from 0.01 to 1.

15. The process according to claim 14, wherein the ratio by weight between the light olefin and the oxygenated compound is in the range of from 0.05 to 0.5.

16. The process according to claims from 1 to 9 and 11 and 14 wherein the oxygenated compound is selected from methanol, dimethyl ether, ethanol or 1-butanol as single or in combination with each others.

17. The process according to claim 16 wherein mixtures of methanol and dimethyl ether contain at least the 50%w. of methanol and/or dimethyl ether.

18. The process according to claims from 1 to 7 wherein the pressure into said second tubular catalytic reactor is equal to that one kept into said first tubular catalytic reactor except pressure drops .

19. The process according to claim 10, wherein said catalyst calcination is carried out at a temperature in the range of from 4500C to 5500C, wherein said catalyst calcination is carried out at a pressure in the range of from 1 bar absolute to 3 bar absolute, wherein said catalyst calcination is carried out with mixtures of oxygen and nitrogen wherein the O2 level is in the range of from 0.1% to 20% by volume, wherein said catalyst calcination is carried out with a space velocity (GHSV) is in the range of from 1000 h-1 to 3000 h-1.

20. The process according to claims 19, wherein said catalyst calcination is carried out in the same reactor wherein the catalyst is placed for the reaction.

21. The process according to one of the previous claims wherein the used catalysts comprise zeolites H-ZSM5;

22. The process according to claim 21, wherein the first catalyst consists of zeolites with molar ratio SiO2Ml2O3 > 100.

23. The process according to claim 22, wherein the second catalyst consists of zeolites with molar ratio SiO2/Al2O3 > 20.

24. The process according to claims from 1 to 7 wherein the reaction product mainly containing a mixture of heavy olefins Cn with n ≥ 5 is hydrogenated to form gasoline cuts or is oligomerized to form mixtures of gasoline, kerosene and diesel .

Description:
PROCESS FOR THE DIRECT CONVERSION OF OXYGENATED COMPOUNDS TO LIQUID HYDROCARBONS HAVING A REDUCED AROMATIC CONTENT

The present invention relates to a process for the direct conversion of oxygenated compounds to liquid hydrocarbons with a reduced aromatic content. More specifically, the invention relates to a process for the transformation of alcohols and/or ethers represented by the formula (C n H 2n+ i) -O- (Q n H 2n .+-.) / wherein n=l-4 and m=0-4, such as methanol, dimethyl ether, ethanol or 1-butanol, alone or combined with each other, into liquid hydrocarbons with a reduced aromatic content, suitable for obtaining gasoline cuts, diesel or jet fuel (kerosene) cuts. The fractions obtained have an aromatic content lower than 10% by weight.

A prominent development activity has been recently directed towards defining ways for the production of hydrocarbon fuels for motor vehicles starting from sources alternative to oil (A. Corma et al . Chem. Rev. 2006, 106, 4044-4098) . It is known in the state of the art, for example, that alternative sources such as coal, natural gas or even biomasses, can be transformed by partial oxidation or steam reforming, into synthesis gas, i.e. a mixture consisting of carbon dioxide and hydrogen. The syngas can be transformed into hydrocarbons through the Fischer-Tropsch process, typically in the presence of heterogeneous catalysts based on cobalt or iron. The Fischer-Tropsch process has a high technological complexity, due, among other things, to the wide range of products obtained, and requires elevated investment costs .

Alternatively, the synthesis gas can be transformed with a high yield into methanol and/or dimethyl ether, typically in the presence of zinc, chromium and copper- based catalysts. It is known, in turn, that oxygenated compounds such as methanol and/or dimethyl ether can be converted to gasoline with a high octane number having a high aromatic content, by means of catalytic conversion on zeolites, in particular ZSM-5 zeolites, according to the Methanol to Gasoline (MTG) process described in various patents: US 3 894 103, US 3 894 104, US 3 894 106, US 3 894 107, US 4 035 430, US 4 058 576. More details can be found in scientific literature, in the articles indicated hereunder for illustrative purposes: a. Blauwhoff et al . , Zeolites as catalysts in industrial processes, b. J. Weitkamp, L. Puppe (eds.) Catalysis and zeolites: fundamentals and applications, Chapter 7. The production of gasoline from synthesis gas via methanol, represents an interesting alternative to the technology via the Fischer Tropsch process. The gasoline obtained with the above-mentioned MTG process has the disadvantage of being characterized by an extremely high quantity of aromatics, higher than 30% by weight, whereas the most recent regulations envisage their limitation; in addition, among the aromatics produced, the durene (1, 2, 4, 5-tetramethyl benzene) content appears to be excessively high, undesirable due to .its high melting point, which causes the tendency to separate from gasoline, with consequent problems in its use for motor vehicles. For this reason, the gasoline obtained with the MTG process must be further treated to reduce the content of durene and also other aromatic components, as described, for example, in US 4 304 951, with consequent technical and economical burdens. More generally, through the MTG process it is not possible to obtain kerosene or diesel cuts, for whose production this process cannot therefore represent an alternative to the Fischer Tropsch process .

In order to avoid these limitations, processes have been described in the state of the art which envisage, in a first stage, the intermediate production, from oxygenated products such as methanol and/or dimethyl ether, of light olefins, typically C 2 -C 5 . The olefins are then converted in a second stage to gasoline, kerosene or diesel cuts, by means of oligomerization processes.

Processes of this type are described, for example, in patents: US 4 025 576, US 4 476 338, US 4 482 772, US 4

497 968, US 4 506 106, US 4 543 435, US 4 547 602, US 4

579 999, US 4 689 205, US 4 898 727, US 4 899 002, US 4

929 780, US 5 045 287, US 5 177 279. These processes are also described in the above-mentioned articles: a. Blauwhoff et al . , Zeolites as catalysts in industrial processes, b. J. Weitkamp, L. Puppe (eds.) Catalysis and zeolites: fundamentals and applications, Chapter 7.

These double-stage processes however are complex, in particular with respect to the fact that the effluent coming from the first stage, must be subjected to onerous separation and compression operations before being sent to oligomerization. The effluent of the first stage, in fact, mainly consisting of a mixture of gaseous olefins and water vapour at low pressure, must first be cooled, in order to separate a liquid aqueous phase (possibly- together with a small quantity of a liquid hydrocarbon phase, i.e. gasoline), and subsequently subjected to compression and heating processes in order to be brought to the conditions required for the oligomerization reaction carried out in the second stage. This represents a considerable technical and economical burden. For this reason, it is common opinion that, as the conversion of methanol to diesel with zeolites requires two stages, the Fischer-Tropsch process is preferred for the conversion of synthesis gas to diesel.

US 4 992 611 describes a process for the direct conversion of C n oxygenated compounds with n=l-4 to hydrocarbon fractions of "distillate" cuts (diesel) with a reduced aromatics content. This process which uses a mixed feeding consisting of an oxygenated compound, for example, methanol, and an olefin, for example propylene from refinery off -gas, only partially converts the oxygenated compounds and has not found industrial application. US 6 372 949 describes a process for the conversion, in a single stage, of oxygenated compounds to gasolines and diesel, in the presence of a particular zeolite catalyst. An objective of the present invention is to propose a simple, double stage process, operating at different temperatures and with no intermediate unitary operations such as compression and separation of the stream leaving the first reaction stage, so as to be able to obtain a hydrocarbon stream with a content of aromatics less than 10% by weight. A further objective of the present invention is to produce a hydrocarbon blend allowing gasoline, kerosene and diesel cuts to be obtained. The present invention relates to a double stage process for the direct conversion of oxygenated compounds to liquid hydrocarbons with a low content of aromatics, and effected in two fixed catalytic bed tubular reactors connected in series, wherein said first reactor operates at a higher temperature than the operating temperature of the second reactor, said process comprising the following phases :

feeding said oxygenated compounds comprising alcohols and/or ethers represented by the formula (C n H 2n+I ) -O- (C m H 2m+ i) wherein n=l-4 and m=0-4, to the first reactor, converting said oxygenated compounds, with a first catalyst, to a mixture of hydrocarbons comprising light olefins, typically rich in olefins C n with n=2-4, and water; ■ feeding said mixture of water and hydrocarbons comprising light olefins, coming from said first reactor, to the second reactor, converting it in said second reactor, in the presence of a second catalyst, to a mixture of liquid hydrocarbons which includes heavy olefins C n with n > 5, non- condensable hydrocarbons and water; ■ separating said hydrocarbon mixture into two streams:

■ one stream including a mixture of liquid hydrocarbons and water, which separates into an organic phase and into an aqueous phase, the last eventually recycled to said first and/or second reactor, wherein said liquid hydrocarbons comprise heavy olefins C n with n ≥ 5 and dissolved non-condensable compounds; ■ one stream which includes a mixture of non- condensable hydrocarbons which are recycled to feed said first reactor.

The liquid product obtained, mainly containing a mixture of heavy C n olefins with n ≥ 5 can be hydrogenated to give gasoline cuts or oligomerized according to conventional processes to give mixtures of gasoline, kerosene and diesel.

The process claimed has the main advantage of directly producing blends of liquid hydrocarbons with a low content of aromatics and C n light paraffins with n=l-4 which decrease the process yield to liquid hydrocarbons. In particular, said process allows a liquid product to be obtained, mainly consisting of C n olefins with n≥5 and a quantity of aromatic compounds equal to or lower than about 10% by weight with respect to the weight of the liquid obtained. A further advantage of the process claimed is the technological simplicity as this process does not include any intermediate separation and compression stage. Detailed description

Further objects and advantages of the present invention will appear more evident from the following description. The process proposed can be effected in continuous by- means of two fixed catalytic bed tubular reactors, using two suitably formed catalysts, preferably extruded in the form of pellets, suitable for being used in the reactors adopted, as is well known in the state of the art and as described in more detail hereunder.

The catalysts used in said process are preferably zeolites in acid or at least partially acid form. The catalysts used in said process are more preferably zeolites of the H-ZSM5 type a crystalline, porous, silico-aluminate having an MFI structure. Said catalyst is described in US 3 702 886, in Kokotailo G. T., Lawton S. L., Olson D. H. and Meier W. M. "Structure of synthetic zeolite ZSM-5", Nature, 272, 437-438 (1978), in Olson D. H., Kokotailo G. T., Lawton S. L. and Meier W. M. "Crystal Structure and Structure-Related Properties of ZSM-5", J. Phys. Chem. 85, 2238-2243 (1981) and in the Database of Zeolite Structures - Structure Commission of the International Zeolite Association. The reference to the structure of the ZSM-5 zeolite can also be found in "Atlas of Zeolite Framework Types" Ch. Baerlocher, W.M.Meier and D. H. Olson, Fifth Revised Edition 2001, Elsevier Amsterdam.

The catalyst used in the first reactor can have a crystalline lattice in which the molar ratio (SAR, silica to alumina ratio) between silicon oxide (SiO 2 ) and aluminium oxide (Al 2 O 3 ) is higher than 20, preferably between 25 and 1,000. The zeolite used in the first reactor is more preferably a H-ZSM5 zeolite, with a high SAR ratio, with a SiO 2 /Al 2 O 3 ratio > 100, for example ranging from 100 to 500, even more preferably between 200 and 400, under these conditions, a low formation of aromatic compounds (lower than 10% by weight) is observed and a contemporaneous reduction in the formation C n light paraffins with n=l-4 which lower the yield to liquid hydrocarbons: in this way the selectivity to the desired C n olefins with n≥5, is maximized. The catalyst used in the second reactor can be a zeolite belonging to the family having a MFI structure, preferably a H-ZSM5 zeolite in acid or partially acid form, with a SAR molar ratio preferably lower than that of the zeolite used in the first reactor, normally higher than 20, preferably between 20 and 500.

The use of a catalyst with a high SAR in the first reactor and a catalyst with a low SAR in the second reactor is in any case advisable. The catalyst used in the second reactor can be the same as that used in the first reactor. To enable these catalysts to be used in fixed bed tubular reactors, the porous crystalline zeolite material of which they consist, must be formed by mixing, according to the known techniques, the zeolite, in the form of powder crystal, with a suitable inorganic binder which must be sufficiently inert with respect to reagents and products, for example silica, alumina, clays (bentonite, kaolin) , or other metal oxide materials such as zirconia, magnesia and mixtures thereof. Alumina is the preferred binder and can be introduced into the catalytic composition by means of one of its precursors, such as, for example, bohemite or pseudobohemite, which generates alumina by calcination. In the preparation of the catalytic composition formed, the zeolite is preferably used in the form of ammonia, which is then transformed into the corresponding acid form by calcination. The relative proportions of the porous crystalline material and binder range from 5:95 to 95:5 by weight, more preferably from 20:80 to 80:20 by weight. The zeolite/binder composite material is produced in a form and dimension suitable for being used in an industrial reactor, above all for the purpose of obtaining a low pressure drop and suitable mechanical resistance and abrasion resistance of the material. The catalyst can be prepared according to spherulization, extrusion, pelletting and granulation processes known in the state of the art; extrusion is the preferred process. The extrusion processes include the use of a peptizing agent which is mixed with the zeolite and binder until a homogeneous paste is obtained, ready for the actual extrusion. For the purpose of this invention, cylinders having a diameter of 1-6 mm and a length of 2-20 mm are adequate for the purposes of the invention, but other forms and dimensions can also be used. According to the known techniques, a calcination step follows the extrusion, for example at 550 0 C under an air stream, for 10 hours .

As is known from patent and scientific literature (EP0172686) , it is preferable to subject both catalysts to treatment with water vapour (i.e. steaming) under suitable conditions and before starting the process, by- putting the catalyst in contact with water vapour or gaseous mixtures containing water vapour, in a volumetric concentration ranging from 10 to 100%, more preferably- ranging from 50 to 100%, at a temperature preferably ranging from 300 0 C to 600 0 C, more preferably ranging from 400 0 C to 55O 0 C, and at a pressure preferably ranging from 1 bar absolute and 10 bar absolute, more preferably between 1 bar absolute and 2 bar absolute, preferably for 1-500 hrs, more preferably for 100-500 hrs . The pre- treatment is preferably carried out in the same reactors. The feeding to the first reactor comprises oxygenated compounds including alcohols and/or ethers represented by the formula (C n H 2n+I ) -O- (C m H 2m+ i) wherein n=l-4 and m=0-4.

Methanol, dimethyl ether, ethanol or 1-butanol are preferred, used alone or combined with each other. More preferably the methanol and dimethyl ether blends contain at least 50% by weight of methanol and/or dimethyl ether. The feeding also preferably contains C n light olefins with n=2-4 or blends thereof, deriving, for example from petrochemical or refinery operations, for example from off-gas of fluid bed catalytic cracking (FCC) , or from recycled streams present in the same process, object of the invention. The feeding to the first reactor is preferably diluted with inert gases, such as nitrogen and/or water. In this case, the molar ratio between the inert gas or water, and the oxygenated compound preferably ranges from 0.01 to 10, more preferably from 0.05 to 5. The water is preferably added to the first reactor by means of the recycling coming from the second reactor, which includes C n light olefins with n=2-4. The dilution of the methanol in the feeding with water is amply described in literature by the patent WO 2007/042124. Other components of the hydrocarbon type, such as olefins and paraffins, are preferably added to the feeding of the first reactor.

The reaction temperature in the first reactor preferably ranges from 350 0 C to 600 0 C, more preferably from 400 0 C to 550 0 C. The oxygenated compound is fed to the first reactor at a WHSV (Weight Hourly Space Velocity, expressed as weight of material fed per weight unit of catalyst) preferably ranging from 0.1 h '1 to 100 h "1 , more preferably from 0.5 h "1 to 10 h "1 , even more preferably between 1 h "1 and 5 h "1 . The pressure in the reactor is preferably maintained at 1 bar absolute to 10 bar absolute, more preferably between 1 bar absolute and 5 bar absolute .

By operating in the first reactor under the conditions described, a conversion of the oxygenated compound to hydrocarbons higher than about 95%, more commonly higher than about 99%, is observed, for a lengthy duration of the operating cycle of the first reactor, normally higher than at least 100 hours and which, in relation to the operating conditions adopted, can reach over 500 hours. When the conversion of the oxygenated compound at the outlet of the reactor proves to be lower than the threshold value desired, corresponding, for example, to a value ranging from about 95 to 99%, the operating cycle of the first reactor is interrupted and the catalyst subjected to calcination, even in situ.

The reaction temperature in the second reactor preferably ranges from 250 0 C to 400 0 C, more preferably from 300 0 C to 350 0 C. The pressure in the second reactor is preferably maintained at between 1 bar absolute and 10 bar absolute, more preferably between 1 bar absolute and 5 bar absolute. It is preferable to operate at the same pressure adopted in the first reactor, with the exception of possible pressure drops, or at lower pressure, in order to be able to feed the second reactor directly with the reaction stream leaving the first reactor without the necessity of intermediate compression. The reaction mixture is fed to the second reactor at a WHSV (Weight Hourly Space Velocity, expressed as weight of material fed per weight unit of catalyst) preferably ranging from 0.1 to 100 h "1 , more preferably from 0.5 to 10 h "1 , even more preferably from 1 h '1 to 51T 1 . By operating under the conditions described above, a yield to C n liquid hydrocarbons with n>5 of about 90% by weight, calculated with respect to the weight of the hydrocarbon component of the oxygenated compound, can normally be obtained. Also in this case, when the quantity of C n liquid hydrocarbons with n>5 formed, proves to be lower than the desired threshold value, corresponding, for example, to a yield value per passage ranging from about 50 to 55% by weight, the operating cycle of the second reactor can be interrupted and the catalyst subjected to calcination even in situ.

The calcination is preferably carried out at a temperature ranging from 450 to 550 0 C, at a pressure preferably ranging from 1 bar absolute to 3 bar absolute, with mixtures of oxygen and nitrogen in which the O 2 content preferably ranges from 0.1% to 20% by volume and with a GHSV (Gas Hourly Space Velocity, expressed as litres of gas mixture /hr/litre of catalyst) ranging from 1,000 hr "1 to 3,000 hr "1 .

The calcination is preferably effected in the same reactor as that in which the reaction catalyst is introduced (in situ regeneration) . By operating according to the invention, the catalyst can be regenerated various times, without any modification of the catalytic performances observed.

The effluent from the first reactor contains water and hydrocarbons. Said hydrocarbons generally comprise from 60 to 85% by weight of a mixture of C n light olefins with n=2-4, and for the remaining part C n olefins with n>5 and paraffins. This effluent is fed to the second reactor without adopting any intermediate separation treatment. The quantity of water present in the effluent mixture depends on the feeding to the first reactor and can vary within large ranges, without jeopardizing the catalyst performance.

The stream leaving the second reactor is a mixture which can normally contain: - a liquid mixture comprising:

• between 60 and 85% of C n olefins with n≥5 and water;

• between 15 and 40% by weight of C n light olefins with n=2-4; - a mixture of non-condensable hydrocarbons.

The stream leaving the second reactor, after depressurization and cooling, is subjected to a gas- liquid separation process into two streams:

• a mixture of liquid hydrocarbons and water, said liquid hydrocarbons comprising C n heavy olefins with n ≥ 5 and non-condensable dissolved compounds, and the mixture separates into an organic phase and an aqueous phase possibly recirculated to said first and/or second reactor after possible flushing of the non-condensable dissolved compounds;

• a mixture of non-condensable hydrocarbons which is recirculated to the feeding of said first reactor.

When the light olefins are fed together with the oxygenated compound, the weight ratio between olefin and the oxygenated compound fed preferably ranges from 0.01 to 1, more preferably from 0.05 to 0.5.

The final liquid product obtained mainly consists of C n olefins with ή ≥5 and has a content of aromatics equal to or lower than about 10% by weight with respect to the weight of the liquid obtained.

The final liquid product, mainly containing a mixture of C n heavy olefins with n ≥ 5 is preferably hydrogenated to give gasoline cuts, or preferably oligomerized according to conventional processes, to give mixtures of gasoline, kerosene and diesel.

An embodiment of the invention is inserted, with the help of the enclosed figure 1, provided for purely illustrative and non-limiting purposes. Figure 1 shows a scheme of the process wherein Rl is the first fixed bed, tubular, catalytic reactor and R2 is the second fixed bed, tubular, catalytic reactor, Pl is the pump which charges the reagents, P2 is the pump which recirculates the water, El and E2 are the exchangers which cool the effluents from the reactors Rl and R2 , S is the separator in which the separation takes place, both gas-liquid and aqueous phase and organic phase, streams 2 and 3 are inert products, stream 1 is the feeding of the reagents to Rl, streams 4 and 5 are the effluent from Rl before and after the cooling, streams 6 and 7 are the effluent from R2 before and after the cooling, streams 8, 9 and 10 are the recirculated products containing non-condensable hydrocarbons to Rl and R2 , streams 11 and 12 are the water recirculation to Rl, stream 13 is the final product mainly containing C n heavy olefins with n > 5. EXAMPLE 1

The example is carried out in a micro-pilot plant consisting of two identical tubular reactors (Rl and R2) made of AISI 316L steel, in series, having the following dimensions: height 350 mm, diameter 12,7 mm, volume about 30 ml. Each reactor is equipped with an electric oven which can heat the reactor up to a temperature of 55O 0 C. Upstream of the single reactors (Rl and R2) , gaseous compounds, such as nitrogen (2 and 3) can be fed, with regulation of the flow-rate by means of a TMF (Thermal Mass Flowmeter) , in addition to liquid components (1 and 12) by means of piston dosing pumps (Pl and P2 pumps for HPCL) . The actual quantity of liquid injected is controlled by means of scales. The reaction temperature is measured in the two reactors at different heights in the catalytic bed with a sliding thermocouple. The plant can operate up to a pressure of 10 bar absolute; once the operating pressure has been established, a regulation valve maintains the pre- fixed pressure. Downstream of the regulation pressure valve (ΔP) , the reaction effluents (6) are cooled by means of water exchangers (E2) to a temperature of 15-20 0 C. The condensable reaction products, i.e. the C n hydrocarbons with n ≥ 5 (13) and water (11) , are condensed and collected for analysis, whereas the non- condensable products (8) are analyzed on line via gas chromatography. The process balances are effected by collecting the liquid effluents for 1-2 hours, separation of the aqueous phase from the organic phase and gas chromatographic analysis of the two separate phases. The light gaseous components are, on the contrary, analyzed in continuous, as already mentioned, approximately every hour. 5 g of a catalyst based on the extruded commercial zeolite H-ZSM5 CBV 28014 CY 1.6 of ZEOLYST (80% zeolite,

20% alumina as binder) , are charged into the first reactor (Rl) . The active phase commercial CBV 28014 ZEOLYST has a molar ratio (SAR, silica to alumina ratio) between silicon oxide (SiO 2 ) and aluminium oxide (Al 2 O 3 ) equal to 280 and a surface area equal to about 400 m 2 /g.

As indicated by the producer

[htpp: //www. zeolyst.com/html/zsm5.asp] , the catalyst is ground and sieved, recovering and introducing into the reactor the fraction ranging from 0.8 to 1.0 mm in diameter. 12 g of corundum are charged below the layer of catalyst, a further 9 g of corundum are charged above (particle size of corundum 0.8-1 mm in diameter) . The pressure of the control valve is set at 1 bar absolute and a flow of 10 Nl/h of nitrogen (2) is fed to the first reactor (Rl) , with the second reactor (R2) excluded from the circuit. The temperature of the first reactor (Rl) is brought to 500 0 C, once the established temperature has been reached, the nitrogen flow (2) is interrupted and the feeding of 15 g/h of water is initiated. The steaming treatment of the catalyst is prolonged for 250 hours. At the end of the treatment, the water feeding is interrupted, a flow of 10 Nl/h of nitrogen (2) is sent for 10 h, the system is then left to cool still under a nitrogen flow (2) . 5 g of the same catalyst based on the commercial zeolite H-ZSM5 CBV 28014 CY 1,6 of ZEOLYST are charged into the second reactor (R2) . The catalyst of the second reactor is used as such, i.e. it is not subjected to the pretreatment with vapour described above for the catalyst of the first reactor. The value of the pressure control valve is set at 3 bar absolute and the feeding of a flow of 10 Nl/h of nitrogen (2) to the first reactor

(Rl) is initiated, the temperature of the first reactor

(Rl) is brought to 460 0 C and the temperature of the second reactor (R2) to 320 0 C. After about 4 hours, the system has reached the temperatures established. At this point the nitrogen feeding (2) is closed and a mixture

(1) of 15 g/h consisting of: 86% w methanol, 9% w isopropanol, 5% w water, is fed to the first reactor

(Rl) . The isopropanol, as precursor of the C 3 olefin, simulates the recycling of light olefins in the process.

Table 1 shows the results at different sampling times.

Table 1

(*) DME (dimethyl ether < 0.1% w) . EXAMPLE 2

At the end of the test described in EXAMPLE 1, the value of the pressure control valve is set at 1 bar absolute, the feeding to the plant of a flow of nitrogen (2 and 3) equal to 10 Nl/h (2 and 3) is initiated, to both the first and second reactor (Rl and R2) , and at the same time the temperature of both reactors is brought to 500 0 C. Once the desired temperature has been reached in the two reactors, flushing with nitrogen is continued for 2 h, at this point a mixture of 10 Nl/h of nitrogen/air (2 and 3) is fed at a ratio of 80/20 by volume for about 2O h. In order to complete the regeneration, a mixture of 10 Nl/h of nitrogen/air is fed at a ratio of 20/80 by volume for an additional 4 hours. The experimentation is repeated under the same conditions as described in Example 1 up to about 100 hours of reaction, the regeneration phase is then repeated as described above. Reaction and regeneration phases are alternated until a total of about 500 hours of reaction are reached, corresponding to 5 reaction/regeneration cycles. Table 2 indicates the performances of the catalyst after the 5 th regeneration .

Table 2

(*) DME (dimethyl ether < 0.2% w) . EXAMPLE 3

Reaction and regeneration phases are alternated as described in the previous example until a total of about 1,000 hours of reaction are reached, corresponding to 10 reaction/regeneration cycles. Table 3 indicates the catalyst performances after the 10 th regeneration.

Table 3

(*) DME (dimethyl ether < 0.2% w) . EXAMPLE 4

Example 1 is repeated with the only difference that the pressure of the regulation valve is set at 2 bar absolute. Table 4 indicates the results at different sampling times.

Table 4

(*) DME (dimethyl ether < 0.5% w) EXAMPLE 5

At the end of the test described in Example 4, the catalyst is regenerated as described in Example 2. The experimentation is then repeated under the same conditions described in Example 4. Reaction and regeneration phases are alternated until a total of about 1,000 hours of reaction are reached, corresponding to 10 reaction/regeneration cycles. Table 5 indicates the process performances after the 10 th regeneration.

Table 5

(*) DME (dimethyl ether < 0.5% w) . EXAMPLE 6

In order to study the effect of a recycling of C 3 -C 4 olefins to the second reactor, the following experiment is effected. The catalysts used in Example 5 are regenerated according to the procedure described in Example 2. The experimentation is then repeated under the same conditions described in Example 4 with the difference that in the feeding to the second reactor, 2.6 g/h of a mixture of C 3 -C 4 olefins (40%w propylene, 20%w 1-butene, 20%w isobutene, 20%w 2-butenes cis-trans) and 3.3 g/h of water are added to the flow leaving the first reactor (4) . Table 6 indicates the process performances.

Table 6

(*) DME (dimethyl ether < 0.5% w) .




 
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