Title of Invention

"AN IMPROVED PROCESS FOR THE CONVERSION OF C1-2 ALKANE OR MIXTURE OF THE SAID ALKANES OR A FEED CONTAINING SAID ALKANES (S)TO AROMATICS"

Abstract An improved process for the conversion of C1-2 alkane or a mixture of the said alkanes or a feed containing said alkane(s) to aromatics by treating a bifunctional pentasil zeolyte catalyst optionally containing one or more transition element having strong dehydrogenation and acid sites with a mixture of H2, steam and in the presence of inert gas at a gas hourly space velocity of at least 500 cm3 g-1 h-1 at a temperature in the range of 400°C-800°C and pressure in the range of 1-5 atm for a period of at least 0.5 h, treating the catalyst obtained in step (i) with air or O2 at a gas hourly space velocity of at least about 200 cm3 g-1 h-1 at a temperature in the range of 400°C-800°C and pressure in the range of 1-5 atm for a period of at least 0.2 h and contacting the catalyst obtained in step (ii) with a lower alkane or mixture of lower alkanes and at least one olefin or at least one higher paraffin or both, at a gas hourly space velocity in the range of 1000-100000 cm3 g-1 h-1, at a temperature in the range of 300°C-600°C and pressure in the range of 1-5 atm, thereby forming a reaction mixture containing aromatics, separating the aromatics formed from the reaction mixture by known methods and if desired and recycling the unconverted lower alkanes and non- aromatics to aromatics.
Full Text FIELD OF THE INVENTION
This invention relates to an improved process for the conversion of C1-2 alkane or a mixture of the said alkanes or a feed containing said alkane(s) to aromatics. This invention particularly relates to a improved catalytic process for the direct conversion of methane or ethane or a mixture thereof or hydrocarbon feed containing such alkanes to aromatics or higher hydrocarbons at low temperature t employing, at least one olefin and/or at least one higher paraffin in the feed, in the presence of bifunctional zeolite catalyst having dehydrogenation and acid functions. The process of the preser invention could be used in petroleum and petrochemical industries for producing aromatic hydrocarbon from a feed stock comprising lower lakane or mixture of lower alkanes such as natural gas and olefin: and/or higher paraffins.
BACKGROUND OF THE INVENTION
Methane is the major constituent of natural gas and also of biogas. World reserves of natural gas are constantly being upgraded and more natural gas being discovered than oil. Because of the problems associated with transportation of a very large volumes of natural gas, most of the natural gas produced along with oil, particularly, at remote places, is flared and hence wasted. The conversion of alkanes contained in the natural gas directly to higher hydrocarbons and aromatics is extremely difficult. If technologies are made available for the conversion of the natural gas to easily transportable less volatile value added products such as aromatic hydrocarbons, a far reaching economic impact can be achieved which will also lead to exploration of more gas-rich field increasing the natural gas reserves.
Aromatic hydrocarbons are important commodity chemicals in the petroleum and petrochemical industries. The most commercially important aromatics are benzene, toluene, ethylbenzene and xylenes. Aromatics are currently produced by catalytic reforming of various petroleum feed stocks and catalytic cracking of naphthas. Aromatics can also be produced by catalytic conversion of alcohols (particularly methanol), olefins or lower alkanes (particularly propane, butanes or LPG). The catalyst used in these processes (methanol-to-gasoline Mobil's MTG process, olefins-to-gasoline-and-distillate or MOGD and M2 forming, both developed by Mobil Oil, and LPG-to-aromatics conversion process or Cyclar Process developed by UOP) belong to the pentasil zeolite family, particularly that having the ZSM-5 structure.
Oxidative Activation of Methane for its Conversion to Aromatics
An oxidative activation of methane for converting it directly to C2 - hydrocarbons, ethane and ethylene, is known in the prior art and it is described in a book [E.E.Wolf "Methane Conversion by Oxidative Process : Fundamental and Engineering Aspects" Van Nostrand Trinhold Catalysis Series, New York, (1992)] and also in a number of review articles [J.R.Aderson, Appl.Caial., 47 (1989) 177; J.S.Lee et. al., Catal.Rev.-Sci.Eng., 39 (1988) 249; G.J.Hutchings et. al., Chem. Soc. Rev., 18 (1989) 25; J.H.Lunsford, Catal.Today 6 (1990) 235; J.HXunsford, Angew.Chem.Intl.Ed.Engl. 34 (1995) 970].
According to a recent U.S. Patent 5,336,825 (1994) of Choudhary V.R. and coworkers, methane can be converted to gasoline range hydrocarbons comprising aromatics hydrocarbons by carrying out the conversion of methane in the following two steps. Step (i) : Catalytic oxidative conversion of methane to ethylene and minor amounts of C3 and C4 olefins in presence of free oxygen using basic solid catalyst at a temperature preferably between 600°C and 850°C. Step (i) : Catalytic conversion of ethylene and higher olefins formed in the step (i) to liquid hydrocarbons of gasoline range over acidic solid catalyst containing high silica pentasil zeolite, using product stream of the step (i) as the feed.
In the other multisteps process, described in Eur, PaLAppl. EP 516,507 (1992) and Fr. Appl. 91/6,195 (1991), a methane rich fraction of natural gas is selectivity oxidized, mixed with the C2+ hydrocarbons rich natural gas fraction and pyrolized and then the mixture is aromatized using an aromatization catalyst based on zeolite and gallium.
All the prior art processes based on the oxidative activation of methane involve
following undesirable highly exothermic methane combustion reactions :
CH4 + 202 → CO2 + 2H2O (1)
CH4+ 1.502 → CO + 2H2O (2)
Hence, these processes are hazardous in nature. Moreover, in these processes, undesirable carbon oxides, CO and CO2, are formed thus reducing the product selectivity and also creating environmental pollution problems.
Non-oxidative Activation of Methane for its Conversion to Aromatics
High temperature non-oxidation of methane for its conversion to C2+ hydrocarbons is known in the prior art.
It has been long known that methane and natural gas can be pyrolytically converted to benzene at temperatures above 899°C (1659°F), preferably above 1200°C.
A paper on "High Temperature Synthesis of Aromatics Hydrocarbons from Methane" published in Science 153 (1966) 1393 disclosed that aromatic hydrocarbons can be prepared from methane by contact with silica at 1000°C. The yield of hydrocarbons was in the range of 4.8 - 7.2 % based on the methane used in the single pass at a gas space velocity of 1224 h" .
More recent, a non-oxidative activation of methane by its dehydrogenase coupling of methane over active carbon at temperature ≥ 1100°C has been reported by H.Yagita et. al. [ Ref. H. Yagita et. al., in Environmental Catalysis, G.Centi et. al. Eds. SCI Publication, Rome, 1995, page 639 - 642].
A U.S. Patent 4,814,533 (1989) discloses a contineous catalytic process for the production of higher molecular weight hydrocarbons rich in ethylene or aromatics or both from lower molecular weight hydrocarbons or methane in which a lower molecular weight hydrocarbon containing a gas is contacted in a reaction zone with a higher molecular weight hydrocarbons synthesis catalyst at a temperature greater than 1OO0°C.
A recent Japanese patent, Jpn. Kokai Tokkyo Koho JP. 07,155,600 (1995), discloses a process for the preparation of reaction media for aromatization and preparation of aromatics from methane using media at high temperature. A reactor containing the reaction media, which is prepared by thermal decomposition of cyclohexane at 1050°C, was feed with methane at 1050°C for 2 h to give benzene in 54.7 % selectivity at 30.9 % conversion.
Because of the requirement of high temperature for the conversion of methane and also due to the extensive coke formation at the high reaction temperature, the above processes based on the non-oxidative conversion of methane are difficult to practice and hence uneconomical.
Catalytic aromatization of methane in the absence of O2 using zeolite catalyst is also known in the prior art.
U.S. Patent 4,727,206, GB Patent 8531687 and European Patent Application No. 0 228 267 Al discloses aromatization of methane by contacting with gallium loaded zeolite containing group VII B metal or metal compound as catalyst at a temperature from 600°C
to 800°C, preferably from 650°C to 775°C, in the absence of oxygen. However, the conversion of methane into aromatics and the yield of the aromatics found reported in the examples of these patents, are very low. At a weight hourly space velocity of 1.0 and absolute pressure of 7.0 bar, the methane conversion at 675°C, 700°C and 750°C was 3.6 wt. %, 4.9 wt. % and 8.3 wt. %, respectively and aromatics yield was 2.0 wt. %, 2.53 wt. % and 2.95 wt. %, respectively. Because of the very low aromatics yield even at a temperature as high as 750°C, this process can not be practically economical.
A U.S. Patent 5,026,937 (1991) discloses a process for the aromatization of methane using a catalyst comprising about 0.1 to about 2 wt. % gallium containing ZSM zeolite and phosphorus-containing alumna, at a gas hourly space velocity of 400 - 7500 h"1 at relatively low temperature conditions. As per the illustrated example of this process, when the catalyst was contacted with a stream of 100 mole % methane at a flow rate of 1.4 h'1 LHSV (liquid hourly space velocity) at 750°C and at atmospheric pressure, the overall methane conversion was 3.5 mole % in the single pass, the selectivity to C2+ hydrocarbons was 72 %, and the selectivity to coke was 28 %. Because of the very low methane conversion even at 750°C and low space velocity and also due to the very high selectivity to coke, this process is not economical. Because of the extensive coke formation, this process is also difficult to practice on a commercial basis.
Although aromatization of methane at or below 600°C temperature is desirable for making the conversion of methane into aromatics process commercially more feasible, the aromatization of methane alone at the low temperatures is not at all thermodynamically possible. At or below 600°C temperature, the conversion of methane to benzene, according to following reaction,
(Formula Removed)
involves a very large free energy change, AGt. The value of AGt, which is greater than 48 kcal per mole of benzene formed at or below 600°C, is much larger than zero. This high thermodynamic barrier does not allow the formation of benzene from methane at the lower temperatures. Hence, for the conversion of methane to aromatics at or below 600°C, it is necessary to find ways for overcoming the thermodynamic barrier and also for the non-oxidative conversion of methane, which is most inert among hydrocarbons, at the lower temperatures.
Activation of Ethane for its Conversion to Aromatics
Ethane is the minor constituent of natural gas. Ethane can be produced by the oxidative coupling of methane which is main constituent of natural gas. Ethane is also formed in the Cyclar LPG aromatization process developed by UOP, as an undesirable byproduct to an appreciable extent [ Ref. M.Guisnet et. al. Appl. Catal. A.Gen. 89 (1992) 1-30 (a review)]. The conversion of ethane to aromatics is therefore, of great practical importance. A few process for the aromatization of ethane are known in the prior art.
A U.S. Patent by Chu (US 4,120,910) disclosed a process for converting ethane to liquid aromatics in the presence of crystalline aluminosilicate zeolite containing catalyst. According to this process, ethane at 1100°F (593.3°C) and 1 aim was passed over a Cu-Zn H-ZSM-5 zeolite (prepared by ion exchange of a synthetic NHj-ZSM-5 with Cu-nitrate), at space velocity = 500 cm3 g'1 h"2 at STP) to give 31.75 % conversion with aromatics yield of 19.05 % and aromatics selectivity of 60 %.
A U.S. Patent by Chester and Chu (US 4,350,835) taught preparation of aromatic hydrocarbons by passing catalyst containing small amount of Ga. According to this patent, passing ethane over a mixed H-ZSM-5 and alumina catalyst at 1090°F (588°C) at 1 atm and space velocity 0.5 h"J (estimated gas hourly space velocity = about 373 cm3 g"1 h1) for 2 h gave a 21 % conversion of ethane with selectivity for aromatics of 56 % and yield for aromatics of 11.8 %.
A U.S. Patent by Desmond and Henry (US 4,766,265) also disclosed conversion of ethane to aromatic hydrocarbons in the presence of Ga/ZSM-5 or ZSM-11 zeolite catalyst promoted with Re and Ni, Pd, Pt, Rh or Ir at 500° - 700°C. According to this patent, passing ethane over a Ga-exchanged H-ZSM-5 zeolite containing 0.6 % Re and 0.3 % Rh at 640°C and space velocity of 0.73 h'1 (estimated gas hourly space velocity = about 545 cm5 g"1 h l) gave 48.3 % ethane conversion and 60 % and 29 % selectivity and yield, respectively, for the hydrocarbon containing 6 or more carbon atoms.
A European Patent by Bennett and Hall (Eur. Pat. Appl. EP 202,000) also disclosed a process for the aromatization of ethane over 0.1 - 10 % Ga loaded zeolite. According to this process, ethane was passed over a Ga-impregnated MFI zeolite at 625°C, 4.5 bar and space velocity of 1.0 h"1 (estimated gas hourly space velocity = about 747 cm3 g"f h'1) to give 3 7 % ethane conversion, 20 % yield of aromatics and 54 % selectivity for aromatics.
The main disadvantages of the prior art processes for the conversion of ethane to aromatics are as follows :
i) The prior art ethane aromatization processes operate at high temperature, generally
above 600°C. At lower temperature, the ethane conversion and aromatics selectivity
or yield are poor, making the process uneconomical, ii) The prior art ethane aromatization process operate not only at high temperature
(generally above 600°C) but also at low gas hourly space velocity (generally at less
than 1000 cm3 g-1 h-1) and consequently, the productivity of aromatics is low. iii)The catalyst used in the prior art ethane aromatization processes is less active and
selective.
In order to make ethane-to-aromatics conversion commercially attractive and economically feasible, it is necessary to develop a novel process that operates at low temperature (at <. preferably at about and high gas hourly space velocity cm3 g h above cmj using a highly active selective catalyst.> Although aromatization of ethane at low temperature (below 600°C) is desirable for making the ethane aromatization process commercially more feasible, the aromatization of ehtane at the low temperature is severly limited by the reaction thermodynamics. For example at 500°C the conversion of ethane to benzene, according to following reaction,
(Formula Removed)
involved a large free energy change, Gr. The value of Gr for the reaction at 500°C is 7.2 Kcal mole-1. At 400°C, the value of G, is still much high (17.6 Kcal mole-1). Since the Gr values at the low temperatures are much larger than zero, there is a high thermodynamic barrier for the aromatization reaction. The thermodynamic barrier does not favour the formation of benzene from ethane at the lower temperature. Hence, for the aromatization of ehtane at below 600°C, it is necessary to find ways for overcoming the thermodynamic barrier and also for activating ethane at the low temperartures.
The main objective of the present invention is therefore to provide an improved process for the conversion of lower alkane or mixture of lower alkanes such as methane, ethane or their mixtures into aromatics or higher hydrocarbons under non-oxidative conditions and at a low temperature.
Another objective of the present invention is to provide an improved process for the conversion of lower alkane or mixture of lower alkanes at a temperature between 300° C to 600°C at ahigh gas hourly space velocity of 1000-100,000 cm3.g-1h-1 overcoming thermodynamic barrier.
Yet another objective of the present invention is to provide an improved process for the conversion of lower alkane or a mixture of lower alkanes to aromatics with high conversion (above 20%), high seiectivity(above 70%), and high productivity(above 5g (aromatics) g-1 (catalyst) h-1)..
The process of the present invention has been developed on the basis of our findings that in the presense of pretreated bifunctional zeolite catalyst having debydrogenation and acid functions and in the presence of atleast one olefin and / or atleast one higher paraffin at a space hourly velocity in the range of 1000 to 100000 cm3.g-1. h-1 the conversion of the lower alkanes into aromatics or higher hydrocarbons takes place at a temperature in the range of 300-600•° C, overcoming thermodynamic barrier. The conversion under the above said conditions is found to be very high(above 20%) with high selectively(above 70%).
The conversion of methane and/or ethane at low temperature in the process of the present invention can be explained by the hydrogen-transfer and / or alkylation reactions between methane or ethane or both and the olefin over the bifunctional zeolite, involving following elementary reaction steps.
The olefin interacts with zeolitic proton to form a carbenium ion.
(Formula Removed)
where ≥2,H+ = zeolitic proton or protonic acid site and (CnH2n-n)+ is a carbenium ion.
The methane and ethane molecule is partially activated on the non-framework Ga-oxide
species, as follows:
(Formula Removed)
The partially activated methane / ethane molecule, CH3+8—H-δ, interacts with the carbenium ion to form a carbonium ion, which is a pentacoordinated carbocation, as follows :
(Formula Removed)
The carbonium ion, which is an ionic reaction intermediate, is decomposed either to CEb+
carbenium ion and CnH2n+2, involving hydrogen transfer reaction between methane / ethane
and the carbenium ion or to CH3-CnH2a+2 with a release of proton, H+, involving
alkylation reaction between methane and carbenium ion, as follows :
(Formula Removed)
The carbocations : carbenium ions, which are trivalent classical cations, and carbonium
ions, which are non-classical penta or tetracoordinated cations, are described well in the
chemistry literature [ Ref. G.A. Olah, Carbocations and Electrophilic Reactions, Verlag
Chemie, John Wiley & Sons, 1974; Angew.Chem.Int.Ed.vol. 34, page 1393, 1995].
The methylinium ion, H3C+, formed in reaction 8 is rapidly decomposed and releases the proton and CH2 radical; the latter is rapidly dimerized to ethene.
(Formula Removed)
Because of its very high reactivity, the higher paraffin formed in reaction 8 and 9 undergoes fast dehydrogenation over the bifunctional zeolite, converting it to olefin.
(Formula Removed)
The ethene and higher olefins are oligomerized and then dehydrocyclized to aromatics
over the bifunctional zeolite catalyst, as follows :
(Formula Removed)
The methane activation due to the presence of higher paraffin also involves the above reactions but the paraffin is first converted to olefin by reaction 12
Because of the involvement of very high free energy change, AGr > 48 kcal/mole of benzene, according to reaction 3, the direct formation of benzene from methane at £ 600QC is not possible thermodynamics!iy. The thermodynamic barrier is, however, drastically reduced or even eliminated because of the addition of olefins or higher paraffins; the value of AGr approaches to zero or even becomes negative, depending upon the additive , its concentration relative to that of methane, and temperature.
The low temperature activation of ethane occurs because of hydrogen-transfer and / or alkylation reactions between ethane and olefins over the bifunctional zeolite, involving following elementary reaction steps.
The olefin interacts with zeolitic proton to form a carbenium ion.
(Formula Removed)
where n ≥ 2, H+ = zeolitic proton or protonic acid site and (CnH2tt+1)+ is a carbenium ion. Ethane molecule is partially activated on the non-framework Ga-oxide species, as follows:
(Formula Removed)
The partially activated ethane molecule, CsHs—H"5, interacts with the carbenium ion to form a carbonium ion, which is a pentacoordinated carbocation, as follows :
(Formula Removed)
The carbonium ion, which is an ionic reaction intermediate, is decomposed either to C2H5+ carbenium ion and CnH2n+2, involving hydrogen transfer reaction between ethane and the carbenium ion or to CiHs-CnHin+z with a release of proton, H+, involving alkylation reaction between ethane and carbenium ion, as follows :
Carbonium ion  C2H5+ + CnH2n+2 (17)
Carbonium ion  C2H5-CnH2n+i + H4" (18)
The carbocations : carbenium ions, which are trivalent classical cations, and carbonium ions, which are non-classical penta or tetracoordinated cations, are described well in the chemistry literature [ Ref. G.A. Olah, Carbocations and Eiectrophilic Reactions, Verlag Chemie, John Wiley & Sons, 1974; Angew.Chem.Int.Ed.vol. 34, page 1393, 1995].
The carbenium ion, C2H/, formed in reaction 5 is rapidly decomposed to ethylene with a release of the proton.
(Formula Removed)
Because of its very high reactivity, the higher paraffin formed in reaction 17 (when
n > 2) or 18 undergoes fast dehydrogenation over the bifunctional zeolite, converting it to
olefin.
(Formula Removed)
The ethylene and higher olefins are oligomerized and then dehydrocyclized to aromatics
over the bifunctional zeolite catalyst, as follows :
(Formula Removed)
The ethane activation due to the presence of higher paraffin in the feed also involves the above 14-21 reactions but the paraffin is first converted to olefin by reaction 20.
Because of the involvement of high free energy change, AGr > 7.2 kcal/mole of benzene, according to reaction 1, the direct formation of benzene from ethane at <. is not possible thermodynamically. the thermodynamic barrier however drastically reduced or even eliminated because of presence olefins and higher paraffins in feed value ag approaches to zero becomes negative depending upon reaction temperature present their concentration relative that ethane.> SUMMARY OF THE INVENTION
It has now been found by as due to exclusive reseach work that, by adding at least one olefin or at least one higher paraffin or both to lower alkane or mixture of lower alkanes or a feed of natural gas containing lower alkanes in the presense of a pretreated bifunctional zeolyte catalyst, having strong dehydrogenation and acid sites, the thermodynamic barrier for the aromatization of the lower alkane is overcome and the non-oxidative activation of lower alkanes occurs at or below 600*C temperature.
Accordingly the lower alkanes can be converted to aromatics or higher hydrocarbons with high conversion and selectivity. Simultaneously with such conversion of the other hydrocarbons present in the feed also get converted to aromatics or higher hydrocrbons alkanes at the low temperature. The low temperature non-oxidative conversion of the lower alkanes results from hydrogen transfer and / or alkylation reaction of the lower alkane with the olefin, or from higher paraffin present in the feed, over the catalyst. The term "lower alkanes" means the alkanes containing one or two carbon atoms. The higher hydrocarbons are the hydrocarbons containing more than two carbon atoms. The higher paraffin means the paraffin containing more than two carbon atoms.
The main products of the process of this invention are aromatic hydrocarbons comprising benzene, toluene, ortho-xylene, meta-xylene, para-xylene, ethylbenzene, trimethylbeazenes and other aromatics containing 9 and 10 carbon atoms. The minor products of the process of this invention are olefins and paraffins containing 2-4 carbon atoms, along with traces of C5+ aliphatic hydrocarbons.
Accordingly, the present invention provides an improved process for the conversion of C1-2 alkane or a mixture of the said alkanes or a feed containing said alkane(s) to aromatics which comprises : i) treating a bifunctional pentasil zeolyte catalyst optionally containing one or more transition
element such as herein described having strong dehydrogenation and acid sites with a
mixture of H2, steam and in the presence of inert gas at a gas hourly space velocity of at least
500 cm3 g"1 h"1 at a temperature in the range of 400°C-800°C and pressure in the range of 1-5
atm for a period of at least 0.5 h, ii) treating the catalyst obtained in step (i) with air or 02 at a gas hourly space velocity of at least
about 200 cm3 g"1 h"1 at a temperature in the range of 400°C-800°C and pressure in the range
of 1-5 atm for a period of at least 0.2 h, and lii) contacting the catalyst obtained in step (ii) with a lower alkane or mixture of lower alkanes and
at least one olefin or at least one higher paraffin or both, at a gas hourly space velocity in the
range of 1000-100000 cm3 g-1 h-1, at a temperature in the range of 300°C-600°C and pressure
in the range of 1-5 atm, thereby forming a reaction mixture containing aromatics, iv) separating the aromatics formed from the reaction mixture by known methods and if desired, v) recycling the unconverted lower alkanes and non-aromatics to aromatics.
The bifunctional pentasil zeolite which may be employed in the process of the present invention may have ZSM-5, ZSM-8 or ZSM-11 type crystal structure consisting of a large number of 5-membered oxygen-rings i.e. pentasil rings, which are more stable as compared to other 0-rings. The zeolite with ZSM-5 type structure is the more preferred catalyst.
The ZSM-5, ZSM-8 and ZSM-11 type zeolite structures are all well known in the prior art and have unique shape-selective behavior. Zeolite ZSM-5 is described in greater detail in U.S. Patent 3,702.886. ZSM-11 zeolite is described in U.S. Patent 3,709,979. ZSM-8 zeolite is described in Netherlands Patent 7,014,807 and U.S. Patent 3,700,585. ZSM-5/ZSM-11 intermediate zeolite structure are described in U.S. Patent 4,229,424. The term "Zeolite" used herein is not only for microporous crystalline aluminosilicate but also microporous crystalline galloaluminosilicates and gallosilicates.
The bifunctional pentasil zeolite catalyst used may be preferably selected from the group consisting of Ga-containing ZSM-5 type zeolites such as Ga-impreganated H-ZSM-5, Ga-exchanged H-ZSM-5, H-gallosilicate of ZSM-5 type structure and H-galloaluminosilicate of ZSM-5 type structure. These zeolites can also be prepared by methods known in the prior art.
The bifunctional Ga-containing ZSM-5 type pentasil zeolite used in the process of the present
invention contains tetrahedral aluminium and/or gallium, which are present in the zeolite
framework or lattice, and octahedral gallium, which is not present in the
about at least 500 cm3 g'5 h'3 at a temperature in the range of 400° - 800°C and
pressure in the range of 1 - 5 for a period of at least 0.5 h, ii) treating the catalyst obtained in step (i) with air or O3 at a gas hourly space velocity
of at least 200 cm3 g-1 h-1 at a temperature in the range of 400° - 800°C and pressure in
the range of 1 -5 atm for a period of at least 0.2 h, and iii)contacting the catalyst obtained in step (ii) with a lower alkane or mixture of lower
alkanes and at least one olefin or at least one higher paraffin or both, at a gas hourly
space velocity in the range of 1000 - 100000 cm3 g-1 h-1, at a temperature in the
range of 300°C - 600°C and pressure in the range of 1 - 5 atm, iv) separating the higher alkanes / aromatics formed from the reaction mixture by known
methods and if desired , v) recycling the reaction mixture containing unconverted lower alkanes and non-aromatiacs to step(iii) for further conversion.
The Afunctional pentasil zeolite which may be employed in the process of the present invention may have ZSM-5, ZSM-8 or ZSM-11 type crystal structure consisting of a large number of 5-membered oxygen-rings i.e. pentasil rings, which are more stable as compared to other O-rings. The zeolite with ZSM-5 type structure is the more preferred catalyst.
The ZSM-5, ZSM-8 and ZSM-11 type zeolite structures are all well known in the prior art and have unique shape-selective behavior. Zeolite ZSM-5 is described in greater detail in U.S. Patent 3,702,886. ZSM-11 zeolite is described in U.S.Patent 3,709,979. ZSM-8 zeolite is described in Netherlands Patent 7,014,807 and U.S.Patent 3,700,585 ZSM-5/ZSM-11 intermediate zeolite structure are described in U.S.Patent 4,229,424. The terra "Zeolite" used herein is not only for microporous crystalline aluminosilicate but also for microporous crystalline galloaluminosilicates and gallosilicates.
The Afunctional pentasil zeolite catalyst used may be preferably selected from the group consisting of Ga-containing ZSM-5 type zeolites such as Ga-impreganated H-ZSM-5, Ga-exchanged H-ZSM-5, H-gallosilicate of ZSM-5 type structure and H-gaUoaiuminosiiicate of ZSM-5 type structure. These zeolites can also be prepared by methods known in the prior art.
The Afunctional Ga-containing ZSM-5 type pentasil zeolite used in the process of the present invention contains tetrahedral aluminium and / or gallium, which are present in the zeolite framework or lattice, and octahedral gallium, which is not present in the
zeolite framework but present in the zeolite channels in a close vicinity of the zeolitic protonic acid sites, which are attributed to the presence of tetrahedral aluminium and gallium in the zeolite. The tetrahedral or framework Al and / or Ga is responsible for the acid function of the zeolite and the octahedral or non-framework Ga is responsible for the dehydrogenation function of the zeolite.
The most effective and efficient bifunctional pentasil zeolite which can be used in the process of the present invention is H-galloaluminosilicate of ZSM-5 type structure having framework (tetrahedral) Si/Al and Si/Ga mole ratio of about 10 -.100 and 15 - 150, respectively, and non-framework (octahedral) Ga of about 0.5 - 5.0 wt%. These pentasil H-galloaluminosilicate zeolite can be prepared by procedures known in the prior art.
The transition elements if present in the catalyst is selected from Cr, Mo, Fe, Co, Ni, Zn, Re, Ru, Rh, Pd, Os, Ir and Pt. If such elements are present these may be from trace to 10 wt %.
The treatment of the catalyst with the mixture of H2, steam and inert gas if necessary in step (i) is essential. Because of this treatment dehydrogenation in the vicinity of zeolite protonic acid sites in the zeolite are created due to the uniform dispersion of the metal oxide present in the catalyst. The inert gas if used may be selected from N2, He, Ar etc. The inert gas in step (i) is employed for diluting the H2 - steam mixture. The treatment step in (ii) of the catalyst obtained in step (i) with air or Oa is also essential for effecting the dehydrogenation activity. In the treatment step(i), the H2 /inert gas and steam / inert gas ratios may be in the range of 0.05-5.0 and 0.02-2.0, respectively.
In case of the use of pentasil H galloaluminosilicate zeolite, the density of strong acid sites can be controlled by the framework Al/Si mole ratio, higher the Al/Si ratio higher is the density of strong acid sites. The highly dispersed non-framework gallium oxide species can be obtained by the degalliation of the zeolite by its pretreatment with H2 and steam. The zeolite containing strong acid sites with high density and also highly dispersed non-framework gallium oxide species in close proximity of the zeolite acid site is preferred for the process of the present invention. The catalyst may or may not contain any binder such as alumina, silica or clay material. The catalyst can be used in the form of pellets, extrudes and particles of different shapes and sizes.
In the process of this invention, the feed comprising lower alkanes, olefins and/or higher paraffins may be contacted with the catalyst in a single or multiple fixed bed reactors, fluid bed reactor or moving bed reactor, known in the prior art.
The olefins in the feed may be selected from olefins containing 2 - 10 carbon atoms more preferably 2-4 carbon atoms. The preferred higher paraffin in the feed may be selected from paraffins containing 2-10 carbon atoms more preferably 3-6 carbon atoms. The concentration of inert gas N2 if present in the feed may be from traces to 80 mole %; the preferred mole ratio of olefin /or and higher paraffin to methane and/or ethane in the feed ranges from about 0.2 to about 2.0; the preferred gas hourly space velocity of the feed ranges from about 3000 cm3 g"1 h"! to about 50,000 cm3 g"! h'; the preferred temperature ranges from 400°C to 600°C; and the preferred pressure ranges from above 1 atm to about 3 aim.
In a preferred embodiment of the process of this invention, using a feed of comprising 33.3 mole % CH4, 16? mole % iso-butene and 50 mole % N3, the methane and iso-butene present in the feed can be converted to aromatics with 44.2 % and 100%, respectively, conversion and 93.8 % selectivity for aromatics with less than 1% selectivity for coke at 500°C, and gas hourly space velocity of 6200 cm3 g-1 h-1 and pressure of 1.1 atm.
In this specification the various items used convey the following means.
Framework Si means Si present in the lattice of the zeolite.
Framework Al or Ga means the Tetrahedral Al or Ga present in the lattice of the zeolite.
Non-framework Ga means Octahedral Ga present in the zeolite channels.
Gas hourly space velocity, GHSV means volume of feed gases, measured at O°C and 1
atm pressure passed over a unit mass of catalyst per hour.
Conversion, % means weight percent of particular reactant converted to all the products.
Product Selectivity, % = (weight percent of reactant or re act ants converted to a particular
products) x 100.
Conversion given in the examples is per pass conversion.
The present invention is described with reference to the examples given below which are provided to illusrate the invention only and therefore, should not be construed to limit the scope of the invention.
EXAMPLES 1 to 8
These examples illustrate the catalytic process of the present invention for the low temperature non-oxidative conversion of methane and thereby converting it directly to higher hydrocarbons or aromatics using n-butene or iso-butene or propene or ethene or propane or n-hexane as an additive in the feed comprising methane, and using ZSM-5 type H-galloaluminosilicate zeolite having bulk Si/Ga = 24.3, bulk Si/Al = 40.3, framework Si/Ga = 49.9 , framework Si/Al = 40.3, Na/(A1 + Ga) = 0.03, non-framework Ga = 0.32 mmol g-1, crystal morphology or shape = spherical-hexagonal, crystal size = 5.5 ± 1.5 µm and strong acid sites measured in terms of pyridine chemisorbed at 400°C = 0.46 mmol g"3. All the ratio are mole ratios. The zeolite was prepared by the process described in European Patent Application EP 0124271and in the reference : Choudhary et. al. J. Catal. 158 (1996) 23.
A conventional tubular quartz reactor of 12 mm internal diameter packed with the zeolite catalyst of 52 - 72 mesh size particles and kept in the tubular electrical furnace such that the catalyst is in constant temperature zone of the furnace, was used for illustrating the process. The zeolite catalyst packed in the reactor was pretreated in a flow of H2 - steam - N2 mixture with H2/N2 and steam/N2 mole ratio of 2.0 and 0.05, respectively, at a GHSV of 1050 cm3 g"3 h'3 at 550°C for 1 h and then in a flow of air with a GHSV of 1050 cm3 g'3 h-1 at the 550°C temperature for 1 h. The catalytic process is carried out by passing continuously a mixture of methane and N2 with or without hydrocarbon additive, such as n-butene or iso-butene or propene or ethene or propane or n-hexane in the feed over the pretreated zeolite catalyst at different process conditions given in Table 1. The concentration of methane in the feed for all experimental runs was 33.3 mol %. The reactor or reaction temperature was measured by Chromel-Alumel thermocouple located axially in the catalyst bed. The hydrocarbons in the feed and in the products of the reaction were analyzed by an on-line gas chromatograph with a flame ionized detector and computing integrator, using a 3 mm x 3 m Poropak-Q column for separating C1-C4 hydrocarbons and using a 3 mm x 5 m column containing 5% Benton-34 - 5% dinonyphthalate on chromosorb-W for separating aromatic hydrocarbons.
The conversion and aromatics selectivity data were obtained from the feed and product composition, as follows: conversion (%) = [(wt% of reactant in the feed hydrocarbons - wt% of reactant in the product hydrocarbons) / (wt% of reactant in the feed hydrocarbons)] x 100 and aromatic selectivity (%) = [(wt% of the aromatics in the hydrocarbon products) / (100 - wt% of reactants in the hydrocarbon products)] x 100.
The results at different process conditions are presented in Table 1. The conversion of methane and other hydrocarbon reactants coke formed on the catalyst was ≤ 1.0 %.
Table 1 :: Results of the catalytic reactions, Feed = a mixtuxe of
CH4 (33.3 mole %) and N2 with or without olefinic or higher Iparaffinic hydrocarbon additive, A.E?ampie No
(Table Removed)
The results in the Table 1 clearly show that in the absence of olefin or higher paraffin in the feed, the conversion of methane in the process is zero and hence there is no conversion of methane at 600°C and consequently below 600°C. But when an olefin and a higher paraffin is added in the feed, the thermodynamic barrier is overcome and the methane from the feed is activated at 500" - 600°C and thereby it is converted to higher hydrocarbons or aromatics with high conversion and high selectivity for aromatics. The hydrocarbon additive from the feed is also converted to aromatics with very high conversion or even with 100 % conversion.
EXAMPLES - 9 to 19
These examples illustrate the influence of process conditions such as olefin/methane mole ratio in the feed, reaction temperature and gas hourly space velocity, QHSV, on the product distribution, on the conversion of methane and olefin present in the feed and also on the selectivity for aromatics in the process of this invention.
The catalytic process was carried out by the procedures described in EXAMPLES 1 to 8 and also using the same catalyst and reactor described in EXAMPLES 1 to 8, except that the catalyst was preheated in a flow of H2 -steam - N2 mixture with H2/N2 and steam/N2 mole ratio of 0.5 and 0.05, respectively, at a GHSV of 2500 cm3 g"5 h"1 at 600°C for 2 h and then in a flow of air with a QHSV of 2500 cm3 g"1 h"1 at 600°C for 0.5 h, at the process conditions given in Tables 2 and 3. In all the experiments, the concentration of methane in the feed was 33.3 mole % and the feed diluent was N3.
The results showing the influence of hydrocarbon additive, to methane mole ratio in the feed, reaction temperature on the distribution of hydrocarbons in hydrocarbon products, on the conversion of methane and propene or n-butene, and also on the selectivity for aromatics are presented in Table 2.
The results showing the influence of gas hourly space velocity on the performance of the process of the present invention are shown in Table 3.
(Table Removed)
EXAMPLE 20
This example illustrates the catalytic process of the present invention when no feed diluent is used. The catalytic process was carried out by the procedures described in EXAMPLES 1 to 8 and also using same catalyst and reactor described in EXAMPLES 1 to 8, except that the catalyst was pretreated in a flow of H2 - steam - N2 mixture with H2/N2 and steam/N2 mole ratio of 4.0 and 2.0, respectively, at GHSV of 1000 cm3 g-J h ' at 500°C for lb and then in a flow of O2 with GHSV of 500 cms g-l h-1 at 600°C for 5 h, at the following process conditions.
Feed : A mixture of 50 mole % methane and 50 mole % n-butene. n-Butene/CH4. mole ratio : 1.0 GHSV : 13,200 cm3 g-1 h-1 Temperature : 500°C Pressure : 1.1 atm
The results obtained are as follows. Conversion of methane = 31.7 % Conversion of n-butene = 97.2 % Selectivity for aromatics = 87.3 % Productivity for aromatics = 15.6 g Distribution of hydrocarbons in products : 13.6 wt. % CH4, 1.8 wt % C2H4, 0.6 wt. % C2H4, 1.3 wt. % C3H6, 4.5 wt. % CsHs, 2.3 wt. % C4H8, 2.4 wt. % C4H10, 15.1 wt % benzene, 34.7 wt. % toluene, 7.3 wt. % ethylbenzen and para-xyiene, 12.2 wt. % meta-xylene, 0.4 wt. % ortho-xylene and 3.8 wt. % C9 & 10 aromatics.
EXAMPLES 21-29
These examples illustrate the catalytic process of this invention for the low temperature activation of ethane and thereby converting it to higher hydrocarbons or aromatics using olefins and higher paraffins as additive in die feed comprising ethane, and using ZSM-5 type H-galloaluminosilicate zeolite having bulk Si/Ga = 24.3, bulk Si/Al = 40.3, framework Si/Ga = 49.9 , framework Si/Al = 40.3, Na/(Ai + Ga) = 0.03, non-framework Ga = 0.32 mmol g-1 crystal morphology or shape = spherical-
hexagonal, crystal size = 5.5 ± 1.5 µm and strong acid sites measured in terms of pyridine chemisorbed at 400°C = 0.46 mmol g-1. All the ratio are mole ratios. The zeolite was prepared by the process described in European Patent Application EP 0124271and in the reference : Choudhary et. al. J. Catal. 158 (1996) 23.
A conventional tubular quartz reactor of 12 mm internal diameter packed with the zeolite catalyst of 52 - 72 mesh size particles and kept in the tubular electrical furnace such that the catalyst is in constant temperature zone of the furnace, was used for illustrating the process. The zeolite catalyst packed in the reactor was pretreated in a flow of H2 - steam - N2 mixture with H2/N2 and steam/N2 mole ratio of 2.0 and 0.05, respectively, at a GHSV of 1000 cm3 g h-1 at 550°C for 1 h and then in a flow of air with a GHSV of 1050 cms g-1 h-1 at the 550°C temperature for 1 h. The catalytic process is carried out by passing continuously a mixture of ethane and N2 with or without olefinic and higher paraffinic hydrocarbon additives, designated by O and HP, respectively, in the feed over the pretreated zeolite catalyst at different process conditions given in Table 2. The reactor or reaction temperature was measured by Chromel-Alumel thermocouple located ax tally in the catalyst bed. The hydrocarbons in the feed and in the products of the reaction were analyzed by an on-line gas chromatograph with a flame ionized detector and computing integrator, using a 3 mm x 3 m Poropak-Q column for separating C1-C4 hydrocarbons and using a 3 mm x 5 m column containing 5% Benton-34 - 5% dinonyphthalate on chromosorb-W for separating aromatic hydrocarbons.
The conversion and aromatics selectivity data were obtained from the feed and product composition, as follows: conversion (%) = [(wt% of reactant in the feed hydrocarbons - wt% of reactant in the product hydrocarbons) / (wt% of reactant in the feed hydrocarbons)] x 100 and aromatic selectivity (%) = ((wt% of the aromatics in the hydrocarbon products) / (100 - wt% of reactants in the hydrocarbon products)] x 100.
The results in the Table 4 clearly show that in the absence of olefin and higher paraffin in the feed, the conversion of ethane in the process is very small. But when olefin and higher paraffin are added in the feed, the ethane from the feed is activated even at or below 500°C and thereby it is converted to aromatics with high conversion
and high selectivity for aromatics. The hydrocarbon additives from the feed are also converted to aromatics with very high conversion.
Table 4 illustrates that the ethane conversion is increased drastically and also the selectivity for aromatics is increased to a large extent because of the presence of higher paraffins and olefins in the feed. The hydrocarbon product distributions in the nine runs are given in Table 5.
(Table Removed)
Table 5 : Hydrocarbon product distribution in Examples 30-38
(Table Removed)
EXAMPLE - 39
This example illustrates the influence gas hourly space velocity, GHSV, on the product distribution, on the conversion of ethane, n-butylene and n-hexane present in the feed and also on the selectivity for aromatics at 500°C in the process of this invention.
The catalytic process was carried out by the procedures described in EXAMPLES
21-29 and also using the same catalyst and reactor described in EXAMPLES 21-29,
except that the catalyst was pretreated in a flow of H2 - steam - N2 mixture with H2/N2
and steam/N2 mole ratio of 0.5 and 0.05, respectively, at a GHSV of 2500 cm3 g-1 h-1 at
600°C for 2 h and then in a flow of air with a GHSV of 1000 cm3 g-1 h-1 at 600°C for 0.5
h, at the following process conditions for step {iii) :
Feed : A mixture of 33 mole % ethane, 33 mole % of n-butylene, 10 mole % of
n-hexane and balance N2.
GHSV : Varied from 3100 to 80150 cm3 g ' h!.
Pressure : 1.2 ± 0.1 atm.
Temperature : 500°C.
The results at the different gas hourly space velocities are given in Table 6. The results show that even at very high GHSV the conversion of ethane and other hydrocarbons is high and also the selectivity for aromatics is high and consequently the productivity of aromatics is also high.
Table 6 : Results of simultaneous aromatization of ethane, n-butylene and n-hexane.
(Table Removed)
Advantages of the invention :
i) Because of the high thermodynamic barrier, the direct conversion of lower aikane to aromatics or higher hydrocarbons at or below 600°C temperature in the absence of oxygen or other oxidizing agent is not possible thermodynamically. By the process of the present invention, the thermodynamic barrier is overcome and the lower aikane such as methane / ethane can be converted non-oxidatively at or below 600°C temperature to aromatics or higher hydrocarbons or aromatics simultaneously with the aromatization of olefins or higher paraffins added to the feed.
ii) In the process of the invention, the catalyst used for the low temperature non-oxidalive conversion is a bifuctional pentasil zeolite having both acid and dehydrogenation functions which facilitates conversion at low temperature and at non oxidative conditions resulting in higher selectivity for the conversion to aromatics.
iii) By the process of the invention, the lower alkanes and other hydrocarbons present in the feed can be converted to aromatics with high conversion ( above 20% and 80% respectively )and also with very high selectivity (above 70 % )for aromatics and very low selectivity for coke(less than 1%).
iv)The process of the invention is operated not only at the low temperature but also at the high gas hourly space velocity(above 1000 cm3, g1. hl ) and hence the productivity of aromatics is high ( abaove 5g (aromatics). gl (catalyst) . h1 ).
v) By the process of the invention methane which is the most inert hydrocarbons and which is
difficult to activate for direct conversion to hydrocarbons can be controlled to aromatics at low
temperatures.
vi) The process of the invention does not produce undesirable carbon groups and is therefore
great significance and important.




We claim
1. An improved process for the conversion of C1-2 alkane or a mixture of the said alkanes or
a feed containing said alkane(s) to aromatics which comprises :
i) treating a bifunctional pentasil zeolyte catalyst optionally containing one or more transition element such as herein described having strong dehydrogenation and acid sites with a mixture of H2, steam and in the presence of inert gas at a gas hourly space velocity of at least 500 cm3 g"1 h"1 at a temperature in the range of 400°C-800°C and pressure in the range of 1-5 atm for a period of at least 0.5 h,
ii) treating the catalyst obtained in step (i) with air or O2 at a gas hourly space velocity of at least about 200 cm3 g-1 h-1 at a temperature in the range of 400°C-800°C and pressure in the range of 1-5 atm for a period of at least 0.2 h, and
iii) contacting the catalyst obtained in step (ii) with a lower alkane or mixture of lower alkanes and at least one olefin or at least one higher paraffin or both, at a gas hourly space velocity in the range of 1000-100000 cm3 g-1 h-1, at a temperature in the range of 300°C-600°C and pressure in the range of 1-5 atm, thereby forming a reaction mixture containing aromatics,
iv) separating the aromatics formed from the reaction mixture by known methods and if desired,
v) recycling the unconverted lower alkanes and non-aromatics to aromatics.
2. An improved process as claimed in claim 1 wherein the inert gas used in step (i) is
selected from N2, Ar, and He and is present in a concentration from traces to 80 mol%.
3. An improved process as claimed in claims 1-2 wherein the H2/inert gas mole ratio
employed ranges from 0.05-5.0.
4. An improved process as claimed in claims 1-3 wherein the steam/inert gas mole ratio employed ranges from 0.02-2.0.
5. An improved process as claimed in claims 1-4 wherein the feed of step (iii) comprises C2 to C10 olefin.
6. An improved process as claimed in claims 1-5 wherein the feed of step (iii) comprises C2 to Cio paraffin.
7. An improved process as claimed in claims 1-6 wherein the feed of step (iii) is natural gas containing methane and ethane.
8. An improved process as claimed in claims 1-7 wherein the pentasil zeolite catalyst has ZSM-5, ZSM-8, ZSM-11 structure containing large number of 5-member oxygen rings.
9. An improved process as claimed in claims 1-8 wherein the bifunctional pentasil zeolite is selected from a group consisting of Ga-containing ZSM-5 type zeolites Ga-impregnated H-ZSM-5, Ga-exchanged H-ZSM-5, H-gallosilicatae of ZSM-5 type structure and H-galloaluminosilicate of ZSM-5 structure.
10. An improved process as claimed in claims 1-9 wherein the bifunctional pentasil zeolite used is H-galloaluminosilicate of ZSM-5 type structure with framework (tetrahedral) Si/AI and Si/Ga mole ratio of about 10-100 and 15-150 respectively, and non-framework (octahedral) Ga of about 0.5-5.0 wt%.
11. An improved process as claimed in claims 1-10 wherein the transition elements is selected from Cr, Mo, Fe, Co, Ni, Zn, Re, Ru, Rh, Pd, Os, Ir and Pt or a mixture thereof.
12. A process as claimed in claims 1-11 wherein the amount of transition elements present in the catalyst ranges from traces to 10 wt%.
13. An improved process as claimed in claims 1-12 wherein the catalyst contains a binder used is Alumina, Silica and clay.
14. An improved process as claimed in claims 1-13 wherein the mole ratio of olefin or higher paraffin to lower alkane in the feed in step (iii) ranges from about 0.2 to about 2.0.
15. An improved process as claimed in claims 1-14 wherein the gas hourly space velocity of the feed in step (iii) ranges from about 3000 cm3 g"1 h"1 to about 50,000 cm3 g"1 h"1.
16. An improved process as claimed in claims 1-15 wherein the temperature employed in step (III) ranges from 400°C to 600°C.
17. An improved process as claimed in claims 1-16 wherein the pressure employed in step (iii) ranges from 1 atm to 3 atm.
18. An improved process as claimed in claims 1-17 wherein the lower alkane used in the feed of step (iii) is methane or ethane.
19. An improved process for the conversion of C1-2 alkane or a mixture of the said alkanes or a feed containing said alkane(s) to aromatics substantially as herein described with reference to the examples.

Documents:

1139-del-1997-abstract.pdf

1139-del-1997-claims.pdf

1139-del-1997-complete specification (granted).pdf

1139-del-1997-correspondence-others.pdf

1139-del-1997-correspondence-po.pdf

1139-del-1997-description (complete).pdf

1139-del-1997-form-1.pdf

1139-del-1997-form-19.pdf

1139-del-1997-form-2.pdf

1139-del-1997-form-3.pdf

1139-del-1997-form-60.pdf

1139-del-1997-petition-others.pdf


Patent Number 194291
Indian Patent Application Number 1136/DEL/1997
PG Journal Number 41/2004
Publication Date 09-Oct-2004
Grant Date 03-Feb-2006
Date of Filing 02-May-1997
Name of Patentee COUNCIL OF SCIENTIFIC AND INDUSTRAIL RESEARCH
Applicant Address RAFI MARG NEW DELHI-110001,INDIA
Inventors:
# Inventor's Name Inventor's Address
1 MR. TUSHAR VASANT CHOUDHARY NATIONAL CHEMICAL LABRATORY,PUNE INDIA
2 DR.VASANT RAMCHANDRA CHOUDHARY NATIONAL CHEMICAL LABRATORY,PUNE INDIA
3 MR.ANIL KISAN KINAGE NATIONAL CHEMICAL LABRATORY,PUNE INDIA
PCT International Classification Number C07C 2/52
PCT International Application Number N/A
PCT International Filing date
PCT Conventions:
# PCT Application Number Date of Convention Priority Country
1 NA