process for oxidative dehydrogenation of high selectivity and conversion alkanes

专利摘要:
ABSTRACT ?HIGH SELECTIVITY AND CONVERSION ODH PROCESS?. Ethane can be catalytically oxidatively dehydrogenated to ethylene at high conversions and high selectivity in a circulating fluidized bed reactor (CFB) in the presence of oxygen in the feed in an amount above the flammability limit. The reactor has a regeneration reactor attached to regenerate the catalyst and cycle back to the CFB.
公开号:BR112017006648B1
申请号:R112017006648-3
申请日:2015-10-08
公开日:2021-06-08
发明作者:Vasily Simanzhenkov；Shahin Goodarznia；Xiaoliang Gao；Leonid KUSTOV；Aleksey Kucherov；Elena Finashina；Kamal Serhal
申请人:Nova Chemicals (International) S.A.；
IPC主号:

专利说明:

TECHNICAL FIELD
[001] The present invention relates to the oxidative dehydrogenation of lower paraffins in a process of high conversion and high selectivity. Until now, the technique has emphasized that oxidative dehydrogenation reactions must be carried out in a reaction mixture below the lower oxidative combustion limits. As the reaction is “oxygen deficient” the conversion per step tends to be low. However, at the other extreme, one must operate above the upper oxidative combustion limits. Such a process has very short residence times in the reactor, for both the feed stream and the catalyst, and at once provides a high conversion and high selectivity process. Preferably, the reaction is conducted in an apparatus of design for a fluidized bed catalyst cracker. PRIOR TECHNIQUE
[002] The concept of oxidative dehydrogenation of paraffins to olefins (ODH) emerged in the late 1960s. Steam cracking of paraffins was a well-established technology and commercially practiced well before the 1960s. The perceived benefits of ODH are minor operating temperatures which, in turn, reduce greenhouse gas emissions. The downside of ODH processes is the potential for decomposition (decomp). Industrial-scale facilities are expensive and corporations avoid processes that can result in decomposition. As a result, ODH technology was slow to progress.
[003] There are a variety of United States Patents assigned to PetroTex Chemical Corporation issued in the late 1960s that describe the use of various ferrites in a steam cracker to produce olefins from paraffins. The patents include United States Patent Nos. 3,420,911 and 3,420,912 in the names of Woskow et al. The patents teach the introduction of ferrites, such as zinc, cadmium and manganese ferrites (ie, mixed oxides with iron oxide). Ferrites are introduced into a dehydrogenation zone at a temperature of from about 250 °C to about 750 °C at pressures less than 100 psi (689,476 kPa) for a period of less than 2 seconds, typically from 0.005 to 0.9 second. The reaction appears to occur in the presence of steam which can tend to shift the equilibrium in the “wrong” direction. Furthermore, the reaction does not take place in the presence of a catalyst.
[004] Document GB 1,213,181, which appears to correspond in part to previous Petro-Tex patents, reveals that nickel ferrite can be used in the oxidative dehydrogenation process. Reaction conditions are comparable to the aforementioned Petro-Tex patents.
[005] United States Patent No. 6,891,075, issued May 10, 2005 to Liu, assigned to Symyx Technologies, Inc. teaches a catalyst for the oxidative dehydrogenation of a paraffin (alkane) such as ethane. The gaseous feedstock comprises at least alkane and oxygen, but may also include diluents (such as argon, nitrogen etc.) or other components (such as water or carbon dioxide). The dehydrogenation catalyst comprises at least about 2% by weight NiO and a wide range of other elements, preferably Nb, Ta and Co. Although NiO is present in the catalyst, it does not appear to be the source of oxygen for oxidative dehydrogenation of alkane (ethane).
[006] United States Patent No. 6,521,808 issued February 18, 2003 to Ozkan et al., assigned to Ohio State University, teaches sol-gel supported catalysts the oxidative dehydrogenation of ethane to ethylene. The catalyst appears to be a mixed metal system such as Ni-Co-Mo, V-Nb-Mo possibly doped with small amounts of Li, Na, K, Rb and Cs on a mixed silica oxide/titanium oxide support . Again, the catalyst does not supply oxygen for oxidative dehydrogenation, instead gaseous oxygen is included in the feed.
[007] United States Patent No. 4,450,313, issued May 22, 1984 to Eastman et al., assigned to Phillips Petroleum Company, discloses a catalyst of the Li2O-TiO2 composition, which is characterized by a low conversion of ethane not exceeding 10%, despite a high selectivity for ethylene (92%). The main drawback of this catalyst is the high temperature of the oxidative dehydrogenation process, which is close to or higher than 650 °C.
[008] The preparation of a supported catalyst usable for low temperature oxydehydrogenation of ethane to ethylene is disclosed in United States Patent No. 4,596,787 A, issued June 24, 1986, assigned to Union Carboneto Corp. A supported catalyst for the low temperature gas phase oxydehydrogenation of ethane to ethylene is prepared by (a) preparing a precursor solution having soluble and insoluble portions of metal compounds; (B) separation of the soluble portion; (C) impregnating a catalyst support with the soluble portion; and (d) activation of the impregnated support to obtain the catalyst. The calcined catalyst has the composition MoaVbNbcSbdXe. X is nothing or Li, Sc, Na, Be, Mg, Ca, Sr, Ba, Ti, Zr, Hf, Y, Ta, Cr, Fe, Co, Ni, Ce, La Zn Cd Hg Al, Tl, Pb, As, Bi, Te, U, Mn and/or W; a is 0.5-0.9; b is 0.1-0.4; c is 0.001-0.2; d is 0.001-0.1; and is 0.001-0.1 when X is an element. The patent does not teach or suggest a co-comminution of catalyst and support.
[009] Another example of low temperature oxydehydrogenation of ethane to ethylene using a calcined oxide catalyst containing molybdenum, vanadium, niobium and antimony is described in US Pat. 4,524,236 A, issued June 18, 1985, and 4,250,346 A, issued February 10, 1981, both attributed to Union Carboneto Corp. The calcined catalyst contains MoaVbNbcSbdXe in the form of oxides. The catalyst is prepared from a solution of soluble and/or complex compounds and/or compounds of each of the metals. The dry catalyst is calcined by heating to 220-550 °C in air or oxygen. Catalyst precursor solutions can be supported on a support, for example, silica, aluminum oxide, silicon carbide, zirconia, titania or mixtures thereof. The selectivity for ethylene can be greater than 65% for a 50% conversion of ethane.
[010] United States Patent No. 6.624,116, issued September 23, 2003 to Bharadwaj et al. and No. 6,566,573, issued May 20, 2003 to Bharadwaj et al., both assigned to Dow Global Technologies Inc., describe monolithic Pt-Sn-Sb-Cu-Ag systems that were tested in an autothermal regimen at T>750 °C, the starting gas mixture contained hydrogen (H2: O2 = 2:1, GHSV = 180,000 h-1). The catalyst composition is different from that of the present invention and the present invention does not contemplate the use of molecular hydrogen in the feed.
[011] United States Patents No. 4,524,236, issued June 18, 1985 to McCain, assigned to Union Carboneto Corporation and no. 4,899,033 issued February 6, 1990 to Manyik et al., assigned to Union Carbide Chemicals and Plastics Company Inc., discloses V-Mo-Nb-Sb mixed metal oxide catalysts. At 375-400 °C, the conversion of ethane reached 70% with a selectivity close to 71-73%. However, these parameters were only achieved at very low gas hourly space velocities of less than 900 h-1 (ie 720 h-1).
[012] United States Patent No. 7,319,179, granted on January 15, 2008 to Lopez-Nieto et al., attributed to the Superior Council of Scientific Investigations and Polytechnic University of Valencia, discloses Mo-V-Te-Nb-O oxide catalysts that provide a conversion of 50-70% ethane and up to 95% ethylene selectivity (at 38% conversion) at 360-400 °C. The catalysts have the empirical formula MoTehViNbjAkOx, where A is a fifth modifier element. The catalyst is a calcined mixed oxide (at least from Mo, Te, V and Nb), optionally supported on: (i) silica, alumina and/or titania, preferably 20-70% silica by weight of the total supported catalyst; (ii) silicon carbide. The supported catalyst is prepared by conventional methods of precipitation from solutions, drying the precipitate and then calcining.
[013] The preparation of a Mo-Te-V-Nb composition is described in WO 2005058498 A1, published June 30, 2005 (corresponding to United States published application No. 2007149390 A1). Catalyst preparation involves preparing a slurry by combining an inert ceramic carrier with a solution comprising ionic species of Mo, V, Te and Nb, drying the slurry to obtain a particulate product, precalcining the dry product at 150350 ° C in an oxygen-containing atmosphere and calcining the dry product at 350-750 °C under an inert atmosphere. The prepared catalyst exhibits activity and selectivity in the oxidation reaction comparable to the unsupported catalyst.
[014] A process for preparing ethylene from a gaseous feed comprising ethane and oxygen involving contacting the feed with a mixed oxide catalyst containing vanadium, molybdenum, tantalum and tellurium in a reactor to form ethylene effluent is described in the document WO 2006130288 A1, published December 7, 2006, assigned to Celanese Int. Corp. The catalyst has an ethylene selectivity of 50-80%, thus allowing the oxidation of ethane to produce ethylene and acetic acid with high selectivity. The catalyst has the formula Mo1V0.3Ta0.1Te0.3Oz. The catalyst is optionally supported on a support selected from porous silicon dioxide, flamed silicon dioxide, kieselguhr, silica gel, porous and non-porous aluminum oxide, titanium dioxide, zirconium dioxide, thorium dioxide, lanthanum dioxide, oxide of magnesium, calcium oxide, barium oxide, tin oxide, cerium oxide, zinc oxide, boron oxide, boron nitride, boron carbide, boron phosphate, zirconium phosphate, aluminum silicate, silicon nitride , silicon carbide and glass, carbon, carbon fiber, activated carbon, metal oxide or metal mesh and corresponding monoliths; or is encapsulated in a material (preferably silicon dioxide (SiO2), phosphorus pentoxide (P2O5), magnesium oxide (MgO), chromium trioxide (CR2O3), titanium oxide (TiO2), zirconium oxide (ZrO2) or alumina (Al2O3) However, the methods of preparing the supported compositions involve wet chemistry procedures (the solutions are impregnated onto the solid support and then the materials are dried and calcined).
[015] United States Patent No. 5,202,517, issued April 13, 1993 to Minet et al., assigned to Medalert Incorporated, teaches a ceramic tube for use in the conventional dehydrogenation of ethane to ethylene. The “tube” is a ceramic membrane, ethane flows into the tube and hydrogen diffuses out of the tube to improve the reaction kinetics. The reactive ceramic is 5 microns thick on a support 1.5 to 2 mm thick.
[016] United States Patent No. 6,818,189, issued on November 16, 2004 to Adris et al., assigned to SABIC, teaches in passage columns 9 and 10 a process in which the ceramic tiles are wrapped around of a tubular reactor and different reactants flow externally and inside the tube. The patent is directed to the oxidative dehydrogenation of ethane to ethylene.
[017] There is a significant amount of technique in separating ethylene and ethane using silver or copper ions in their +1 oxidation state. See US Patent No. 6,518,476 in Col. 5, lines 10-15, and Col. 16, line 12 - Col. 17, line 57. NOVA Chemicals also described the separation of olefins from non-olefins using ionic liquids (dithiolene in CA 2,415,064 ( now abandoned)). See also U.S. Patent No. 6,120,692 to Exxon; United States Patent No. 6,518,476, issued February 11, 2003 to Union Carbide in columns 16 and 17; the abstract of JP 59172428, published September 29, 1984; and the abstract of JP 59172427, published September 29, 1984.
[018] United States Patent No. 8,107,825, issued September 13, 2011, to Kuznicki et al., assigned to the University of Alberta contains a good sketch of the prior art for separating ethane from ethylene and a method of adsorption using ETS-10.
[019] United States Patent No. 7,411,107 issued August 12, 2008 to Lucy et al., assigned to BP Chemicals Limited, describes a process for the separation of acetic acid from an oxidative dehydrogenation process to convert ethane in ethylene and acetic acid. The process uses a reversible complex of a metal salt (eg Cu or Ag) to separate ethylene (Col. 8). The patent then discloses that acetic acid can be separated from liquids by distillation (Col. 13, lines 35 to 40).
[020] United States Patent Application No. 20110245571, in the name of NOVA Chemicals (International) S.A., teaches the oxidative dehydrogenation of ethane in a fluidized bed in contact with a bed of regenerative oxides to provide oxygen to the reactor. In this process, “free” oxygen is not directly mixed with the raw material, reducing the likelihood of “decomposition”.
[021] United States Patent No. 3,904,703, issued September 9, 1975 to Lo et al., assigned to El Paso Products Company, describes a layered or zoned oxidative reactor, wherein, after a zone for oxidative dehydrogenation, there is an “oxidation zone” after a dehydrogenation zone to oxidize hydrogen to water. After the oxidation zone, there is an adsorption bed to remove water from the reactants before entering a subsequent dehydrogenation zone. This is to reduce the impact of water on downstream dehydrogenation catalysts.
[022] United States Patent Application No. 2010/0256432, published October 7, 2010 in the name of Arnold et al., assigned to Lummus, disclose in paragraphs 86-94 methods for removing residual oxygen from the product stream. A fuel such as hydrogen or a hydrocarbon can be added to a product stream to eliminate residual oxygen. The patent refers to a catalyst, but does not disclose its composition. As mentioned above, it may then be necessary to treat the product stream to eliminate water.
[023] United States Patent Application No. 2004/0010174 (now abandoned) published January 15, 2004 in the name of Wang et al., assigned to ConocoPhillips Company, discloses the use of a circulating fluidized bed reactor ( CFB) (similar in design to an FCC reactor) to conduct an oxidative dehydrogenation. The disclosure teaches in paragraph 40 that the catalyst acts to transport oxygen to the reactor as network oxygen or as adsorbed oxygen. The disclosure teaches about adding air or oxygen to the feed stream.
[024] United States Patent No. 8,519,210, issued August 27, 2013 to Arnold et al., assigned to Lummus Technology Inc., teaches that the oxygen concentration in water may be limited to, with a margin below, the minimum oxygen for combustion, typically including vapor or inert gases to dilute the feed to values below the flammability limits.
[025] The present invention seeks to provide a one-step process to oxidatively dehydrogenate lower paraffins (alkanes, preferably n-alkanes) to produce alpha-olefins. DESCRIPTION OF THE INVENTION
[026] In one embodiment, the present invention provides a process for the oxidative dehydrogenation of one or more alkanes selected from the group consisting of ethane and propane and mixtures thereof in the presence of a supported catalyst selected from the group consisting of:i) catalysts of the formula:VxMoyNbzTemMenOpem Me is a metal selected from the group consisting of Ta, Ti, W, Hf, Zr, Sb and mixtures thereof; x is from 0.1 to 3 provided that when Me is absent, x is greater than 0.5; y is from 0.5 to 1.5; z is from 0.001 to 3; m is from 0.001 to 5; n is from 0 to 2; p is a number to satisfy the valence state of the mixed oxide catalyst; formula: MoaVbNbcTeeOdem which: a is from 0.75 to 1.25, preferably from 0.90 to 1.10; b is from 0.1 to 0.5, preferably from 0.25 to 0.4; c is from 0.1 to 0.5, preferably from 0.1 to 0.35; and is from 0.1 to 0.35 preferably from 0.1 to 0.3; d is a number to satisfy the valence state of the mixed oxide catalyst on a metal oxide support; comprising: a) passing through an oxidative dehydrogenation reactor containing a fluidized bed of said catalyst said one or more alkanes and oxygen at a temperature of 250°C to 450°C, a pressure of 3.447 to 689.47 kPa (0 0.5 to 100 psi) preferably from 103.4 to 344.73 kPa (15 to 50 psi) and a retention time of said one or more alkanes in said reactor of 0.002 to 10 seconds, and to reduce said catalyst, said catalyst having an average retention time in the dehydrogenation reactor of less than 30 seconds; b) feeding said reduced catalyst to a regeneration reactor and passing an air stream optionally with additional nitrogen at a temperature of 250 °C to 400 °C and pressures from 3.447 to 689.47 kPa (0.5 to 100 psi) preferably from 103.4 to 344.73 kPa (15 to 50 psi) across said bed to oxidize said catalyst; and c) passing said oxidized catalyst back to said oxidative dehydrogenation reactor wherein the amount of oxygen in the feed to said reactor is above the upper flammability limit for said feed. The conversion of alkane to alkene is not less than 50% per step and the selectivity for converting alkane to alkene is not less than 0.9.
[027] In the additional modality, the process comprises passing the product stream through one or more scavenging reactors of oxygen. Preferably, the reactors are operated in parallel, and one is being oxidized and the other is being reduced to the lower oxidation state of the metals in the catalyst.
[028] In some embodiments, oxygen sequestering reactors use the same catalyst used in oxidative dehydrogenation reactors.
[029] In a further embodiment, the oxidative dehydrogenation reactor comprises a riser and the regeneration reactor is a separate fluidized bed reactor, said regeneration reactor being connected with said riser to flow the oxidized catalyst back to said riser (eg CFB type reactor [circulating fluidized bed]).
[030] In a further modality, the upper part of said riser comprises a distributor system to improve the temperature control in the reactor [to minimize the combustion of the alkane feed and] to maintain the overall selectivity of the reactor above 90%.
[031] A further embodiment comprises passing one or more of low temperature steam and atomized water in said catalyst stream in said riser to cool the catalyst to control the heat balance of the oxidative dehydrogenation reactor.
[032] In a further embodiment, there is a drop tube between said oxidative dehydrogenation reactor and said regeneration reactor to flow reduced catalyst from said oxidative dehydrogenation reactor to said regeneration reactor.
[033] A further modality comprises passing low temperature steam [countercurrent to the catalyst flow through said downpipe] to remove entrained alkane feed and product.
[034] A further embodiment comprises passing air or a mixture of air and nitrogen through the regeneration reactor in an amount to substantially extract oxygen from the air or a mixture of air and nitrogen and generate a gas product stream comprising not less than 85% by volume of nitrogen.
[035] An additional modality comprises recycling a portion of the reduced effluent stream of oxygen from the regenerator reactor and optionally cooling it and recycling it to the regenerator reactor.
[036] In an additional modality, a CO promoter is added to the regenerator reactor.
[037] A further embodiment comprises separating said alkene product of the oxidative dehydrogenation reactor from water in the product stream of the oxidative dehydrogenation unit.
[038] An additional modality comprises passing unused nitrogen from the effluent stream of the catalyst regeneration reactor to an integrated site unit operation using nitrogen as part of the feedstock.
[039] In a further embodiment, two or more fixed bed reactors are used as scavengers having tubes and valves so that the feed to the fluidized bed oxidative dehydrogenation reactor passes through one or more of the fixed bed reactors having a dehydrogenation which is oxidized to deplete the oxygen catalyst, and passing the product stream through one or more of the fixed bed reactors having an oxygen-depleted oxidative dehydrogenation catalyst, to remove residual oxygen from the product by reaction and bypass flow. product stream to reactors for oxygen depleted reactors and the feed stream flow for oxygen rich reactors.
[040] In a further modality, the integrated site unit operation is selected from an ammonia plant and an acrylonitrile plant, a urea plant and an ammonium nitrate plant.
[041] In an additional modality, the retention time of the catalyst in the oxidative dehydrogenation reactor is less than 30 seconds (preferably less than 10, more desirable less than 5 seconds).
[042] In an additional modality, the retention time of the catalyst in the regeneration reactor is less than 3 minutes.
[043] In a further embodiment, the ratio of catalyst retention time in the regenerator to the catalyst retention time in the oxidative dehydrogenation catalyst is not less than 3.
[044] In a further embodiment, the product stream from the oxidative dehydrogenation reactor and at least a portion of the effluent stream from the regenerator reactor are passed through separate steam generators for heat recovery.
[045] In a further embodiment, the product stream from the oxidative dehydrogenation reactor is cooled and passed through a column to separate alkene combustion products.
[046] In a further embodiment, the product stream from the oxidative dehydrogenation reactor is cooled and passed through an amine unit to remove CO2.
[047] In an additional embodiment, the support is selected from the group consisting of silicon dioxide, fused silicon dioxide, aluminum oxide, titanium dioxide, zirconium dioxide, thorium dioxide, lanthanum oxide, magnesium oxide, calcium oxide, barium oxide, tin oxide, cerium dioxide, zinc oxide, boron oxide, yttrium oxide.
[048] In an additional modality, the alkane is ethane.
[049] In an additional embodiment, the catalyst is of the formula: MoaVbNbcTeeOdem which: a is from 0.90 to 1.10; b is from 0.25 to 0.4; c is from 0.1 to 0.3; and is from 0.1 to 0.3; d is a number to satisfy the valence state of the mixed oxide catalyst on a metal oxide support.
[050] In an additional modality, in the oxidative dehydrogenation reactor, the conversion to ethylene is greater than 60%.
[051] In an additional modality, in the oxidative dehydrogenation reactor, the selectivity for ethylene is greater than 75%. Brief Description of Drawings
[052] Figure 1 is a schematic drawing of a CFB reactor useful according to the present invention.
[053] Figures 2 a and b show the conversion and selectivity of a feed stream comprising ethylene and 25% mol.-% oxygen at a temperature of 355 °C in the presence of a catalyst according to the present invention for a time 60 seconds.
[054] Figures 3 a and b show the conversion and selectivity of a feed stream comprising ethylene and 25% mol.-% oxygen at a temperature of 355 °C in the presence of a catalyst according to the present invention for a time 60 seconds.
[055] Figure 4 is a graph showing time dependence of ethane and O2 conversion (a) and ethylene formation selectivity (b) after switching the gas flow [air to gas mixture] in the Mo-V catalyst -Te-Nb-Ox at 398°C. [2400 h-1]. Dotted lines correspond to equilibrium values.
[056] Figure 5 is a graph showing time dependence of ethane and O2 conversion (a) and ethylene formation selectivity (b) after switching the gas flow [air to gas mixture] in the Mo-V catalyst -Te-Nb-Ox at 398°C. [600 h-1]. Dotted lines correspond to equilibrium values.
[057] Figure 6 is a graph showing time dependence of the selectivity of ethylene formation after switching the gas flow [air to gas mixture] in the Mo-V-Te-Nb-Ox catalyst at 398 °C at different flow rates.
[058] Figure 7 is a graph showing dependence of the conversion of ethane on the amount of ethane supplied to the reactor at different rates after switching the gas flow [air to gas mixture] in the Mo-V-Te-Nb-Ox catalyst at 398°C.
[059] Figure 8 is a graph showing time dependence of conversion of ethane and residual content of O2 (a) and selectivity of ethylene formation (b) after switching the gas flow [O2 to C2H6] in the Mo- catalyst. V-Te-Nb-Ox at 355 °C.
[060] Figure 9 is a graph showing time dependence of conversion of ethane and residual content of O2 (a) and selectivity of ethylene formation (b) after switching the gas flow [O2 to C2H6] in the Mo- catalyst. V-Te-Nb-Ox at 397°C.
[061] Figures 10, 11 and 12 illustrate how a series of three fixed bed catalysts can be used to scavenge oxygen from the product stream in an oxidative dehydrogenation reactor. BEST WAY TO CARRY OUT THE INVENTION Numerical Tracks:
[062] Except in the operational examples or where otherwise indicated, all numbers or expressions referring to amounts of ingredients, reaction conditions, etc. used in the specification and claims are to be considered modified in all cases by the term “about”. Therefore, unless otherwise indicated, the numerical parameters set forth in the following specification and appended claims are approximations which may vary depending on the properties which the present invention seeks to achieve. At the very least, and not as an attempt to limit the application of the equivalents doctrine to the scope of the claims, each numerical parameter should at least be considered in light of the number of significant digits reported and by applying common rounding techniques.
[063] Notwithstanding that the parameters and numerical ranges that establish the broad scope of the invention are approximations, the numerical values established in the specific examples are reported as accurately as possible. Any numerical values, however, inherently contain certain errors necessarily resulting from the standard deviation found in their respective test measurements.
[064] Further, it should be understood that any numerical range recited in this document is intended to include all sub-ranges classified therein. For example, a range of “1 to 10” is intended to include all subranges between and including the minimum recited value of 1 and the maximum recited value of 10; that is, having a minimum value equal to or greater than 1 and a maximum value equal to or less than 10. Because the numerical ranges disclosed are continuous, they include each value between the minimum and maximum values. Unless expressly stated otherwise, the various numerical ranges specified in this order are approximations.
[065] All ranges of compositions expressed herein are limited in total to and do not exceed 100 percent (percent by volume or percent by weight) in practice. When multiple components may be present in a composition, the sum of the maximum amounts of each component may exceed 100 percent, with the understanding that, and as those skilled in the art will readily understand, the actual used amounts of the components will be in accordance with the maximum value of 100 percent. Catalysts:
[066] There are a variety of catalysts that can be used in accordance with the present invention. The following catalyst systems can be used individually or in combination. A person skilled in the art will understand that combinations must be tested on a laboratory scale to determine if there are any antagonistic effects when combinations of catalysts are used.
[067] The oxidative dehydrogenation catalyst of the present invention can be selected from the group consisting of: i) catalysts of the formula:NixAaBbDdOein quex is a number from 0.1 to 0.9, preferably from 0.3 to 0.9 , more preferably from 0.5 to 0.85, more preferably 0.6 to 0.8; a is a number from 0.04 to 0.9; b is a number from 0 to 0.5; d is a number from 0 to 0.0.5;e is a number to satisfy the valence state of the catalyst;A is selected from the group consisting of Ti, Ta, V, Nb, Hf, W, Y, Zn, Zr, Si and Al or mixtures thereof;B is selected from the group consisting of La, Ce, Pr, Nd, Sm, Sb, Sn, Bi, Pb, Tl, In, Te, Cr, Mn, Mo, Fe, Co, Cu , Ru, Rh, Pd, Pt, Ag, Cd, Os, Ir, Au, Hg and mixtures thereof; D is selected from the group consisting of Ca, K, Mg, Li, Na, Sr, Ba, Cs, and Rb and mixtures thereof; eO is oxygen; eii) catalysts of the formula: MofXgYhem queX is selected from the group consisting of Ba, Ca, Cr, Mn, Nb, Ta, Ti, Te, V, W and mixtures thereof; Y is selected from the group consisting of Bi, Ce , Co, Cu, Fe, K, Mg V, Ni, P, Pb, Sb, Si, Sn, Ti, U and mixtures thereof; f = 1; g is 0 to 2; h is 0 to 2, with a provided that the total value of h for Co, Ni, Fe and mixtures thereof is less than 0.5; and catalysts of formula iii) below, and mixtures thereof.
[068] In one embodiment, the catalyst is the catalyst of formula i) where x is from 0.5 to 0.85, a is from 0.15 to 0.5, b is from 0 to 0.1, and d is from 0 to 0.1. In catalyst i), typically A is selected from the group consisting of Ti, Ta, V, Nb, Hf, W, Zr, Si, Al and mixtures thereof, B is selected from the group consisting of La, Ce, Nd, Sb, Sn, Bi, Pb, Cr, Mn, Mo, Fe, Co, Cu, Ru, Rh, Pd, Pt, Ag, Cd, Os, Ir and mixtures thereof and D is selected from the group consisting of Ca, K, Mg, Li, Na, Ba, Cs, Rb and mixtures thereof.
[069] In an alternative embodiment, the catalyst is catalyst ii). In some embodiments of this aspect of the invention, typically X is selected from the group consisting of Ba, Ca, Cr, Mn, Nb, Ti, Te, V, W and mixtures thereof, Y is selected from the group consisting of Bi, Ce , Co, Cu, Fe, K, MgV, Ni, P, Pb, Sb, Sn, Ti and mixtures thereof.
[070] A family of additional particularly useful iii) catalysts iii) comprises one or more catalysts selected from the group consisting of a mixed oxide catalyst of the formula: VxMoyNbzTemMenOp, wherein Me is a metal selected from the group consisting of Ti, Ta , Sb, Hf, W, Y, Zn, Zr, La, Ce, Pr, Nd, Sm, Sn, Bi, Pb Cr, Mn, Fe, Co, Cu, Ru, Rh, Pd, Pt, Ag, Cd, Os, Ir, Au, and mixtures thereof; and x is from 0.1 to 3, preferably from 0.5 to 2.0, more preferably from 0.75 to 1.5 and provided that, when Me is absent, x is greater than 0.5; y is from 0.5 to 1.5, preferably from 0.75 to 1.0; z is from 0.001 to 3, preferably from 0.1 to 2, more preferably from 0.5 to 1.5; m is from 0.001 to 5, preferably from 1 to 4; n is from 0 to 2, preferably n is 0, however, when Me is present, n is preferably from 0.5; at 1.5; ep is a number to satisfy the valence state of the mixed oxide catalyst.
[071] In one embodiment, the catalyst has the formula: MoaVbNbcTeeOdem which: a is from 0.90 to 1.10, preferably 0.95 to 1.1; b is from 0.25 to 0.4, preferably 0. 3 to 0.35; c is from 0.1 to 0.3, preferably 0.1 to 0.15; and is from 0.1 to 0.3, preferably 0.1 to 0.25; and d is a number to satisfy the valence state of the mixed oxide catalyst on a metal oxide support.
[072] In a further modality, in the catalyst, the x:m ratio is from 0.3 to 10, more preferably from 0.5 to 8, desirably from 0.5 to 6.
[073] The methods of preparing the catalysts are known to those skilled in the art.
[074] For example, the catalyst can be prepared by mixing aqueous solutions of soluble metal compounds, such as hydroxides, sulfates, nitrates, halides, salts of lower mono or dicarboxylic acids (C1-5) and ammonium salts or the metal acid per se. For example, the catalyst can be prepared by mixing solutions such as ammonium metavanadate, niobium oxalate, ammonium molybdate, telluric acid etc. The resulting solution is then typically dried in air at 100-150 °C and calcined in a stream of inert gas, such as those selected from the group consisting of N2, He, Ar, Ne and mixtures thereof at 200-600 °C, preferably at 300-500 °C. The calcination step can take from 1 to 20, typically 5 to 15, usually about 10 hours. The resulting oxide is a friable solid typically insoluble in water.
[075] There are several ways in which the oxidative dehydrogenation catalyst can be supported or bound.
[076] Preferred components for forming ceramic supports and for binders include titanium oxides, zirconium, aluminum, magnesium, silicone, phosphates, borophosphate, zirconium phosphate and mixtures thereof, for both fixed and fluidized bed reactors. In the fluidized bed, typically, the catalyst is usually sprayed with the binder, typically forming spherical particles that range in size (effective diameter) from 40-100 µm. However, care must be taken to ensure the particle area is robust enough to minimize friction in the fluid bed.
[077] The support for the catalyst for the fixed bed can also be a ceramic precursor formed from oxides, dioxides, nitrides, carbides selected from the group consisting of silicon dioxide, fused silicon dioxide, aluminum oxide, titanium dioxide, zirconium dioxide, thorium dioxide, lanthanum oxide, magnesium oxide, calcium oxide, barium oxide, tin oxide, cerium dioxide, zinc oxide, boron oxide, boron nitride, boron carbide, oxide yttrium, aluminum silicate, silicon nitride, silicon carbide and mixtures thereof.
[078] In one embodiment, the support for the fixed bed may have a low surface area of less than 20 m2/g, alternatively less than 15 m2/g, alternatively less than 3.0 m2/g for the oxidative dehydrogenation. Such support can be prepared by compression molding. At higher pressures, the interstices within the ceramic precursor being compressed collapse. Depending on the pressure exerted on the support precursor, the surface area of the support can be about 20 to 10 m 2 /g.
[079] Low surface area support can be of any conventional shape, such as spherical, annular, concave, etc.
[080] In the present invention, the catalyst oxidized in the fluidized bed contains one or more matrix oxygen and adsorbed oxygen. The supported catalyst together with added air or preferably oxygen passes together through an oxidative dehydrogenation reactor, and the catalyst is reduced when ethane is converted to ethylene. The supported catalyst is then passed through the downpipe to the regeneration reactor where it is oxidized.
[081] It is important that the support is dried before use (ie, before adding catalyst). Generally, the support can be heated at a temperature of at least 200 °C for up to 24 hours, typically at a temperature of 500 °C to 800 °C for about 2 to 20 hours, preferably 4 to 10 hours. The resulting support will be free of adsorbed water and should have a surface hydroxyl content of about 0.1 to 5 mmoles/g of support, preferably 0.5 to 3 mmoles/g.
[082] The amount of hydroxyl groups on silica can be determined according to the method described by J.B. Peri and A.L. Hensley, Jr., in J. Phys. Chem., 72(8), 2926, 1968, the entire contents of which are incorporated herein by reference.
[083] The dry support for the fixed bed catalyst can then be compressed into the required shape by compression molding. Depending on the particle size of the support, it can be combined with an inert binder to maintain the shape of the compressed part. Shipments:
[084] Typically, the catalyst loading on the support for the fixed bed catalyst provides from 1 to 30% by weight, typically from 5 to 20% by weight, preferably from 8 to 15% by weight of said catalyst and 99 to 70% by weight, typically from 80 to 95% by weight, preferably from 85 to 92% by weight, respectively, of said support.
[085] The catalyst can be added to the support in any variety of ways. For example, the catalyst can be deposited from an aqueous slurry onto one of the low surface area support surfaces by impregnation, wash coating, brushing or spraying. The catalyst may also be coprecipitated from a slurry with the ceramic precursor (e.g., alumina) to form the low surface area supported catalyst.
[086] Catalyst loading for the fluidized bed can be chosen based on a variety of factors including bed volume, alkane flow rate through the bed, energy balance in the bed, binder type, etc. For the fluidized bed, catalyst loading can cover a wide range of values ranging from 10% by weight to 90% by weight, typically above 20% by weight, desirably above 35% by weight.
[087] In the present invention, the feed to the oxidative dehydrogenation reactor includes oxygen in an amount above the upper limit of ignition/explosive. For example, for oxidative dehydrogenation of ethane, typically, oxygen will be present in an amount of not less than about 5% by mol preferably about 18% by mol, eg from about 22 to 27% by mol, or 23 to 26% by mol. It is desirable not to have such a large excess of oxygen as this can reduce selectivity resulting from combustion of feed or end products. Additionally, a very large excess of oxygen in the feed stream may require additional separation steps at the downstream end of the reaction.
[088] The process of the present invention will be described in conjunction with Figure 1 which schematically illustrates a fluidized bed circulation reactor.
[089] In one embodiment, the reactor system 1 comprises a fluidized bed oxidative dehydrogenation reactor 3 and a regenerator reactor 4. The fluidized bed oxidative dehydrogenation reactor riser 2 and the regeneration reactor 4 are connected by a tube of drop 10 leading clean oxidized supported catalyst from regenerator reactor 4 to the riser of oxidative dehydrogenation reactor 2. Each of the riser of the fluidized bed oxidative dehydrogenation reactor 2, the fluidized bed 6 in the dehydrogenation reactor 3 and the regeneration 4 contains fluidized bed of catalyst particles 5, 6 and 7 respectively. In the riser of oxidative dehydrogenation reactor 2 and regeneration reactor 4 above, fluidized catalyst beds 5 and 7 are withdrawal zones 8 and 9, respectively.
[090] The inlet 11 for the downpipe 10 is coupled to the regenerator reactor 4 generally at a point between about 1/3 to 2/3 of the height of the fluidized bed 7. The downpipe 10 enters the bottom of the riser from oxidative dehydrogenation reactor 2. Reactor 2 riser extends to dehydrogenation reactor 3 above fluidized bed fluid level 6 (typically 1/3 to 2/3 height). The reactor riser 2 expands to form an inverted cone disperser 12 to provide a withdrawal zone for the product catalyst. Optionally, a scatter plate 13 can be used above the cone 12. The scatter can have a different shape than an inverted cone; however, care must be taken to ensure substantially uniform gas flow around the disperser.
[091] The oxidative dehydrogenation reactor operates at temperatures below 450 °C, typically from 350° to 450°, pressures from 3.447 to 689.47 kPa (0.5 to 100 psi) preferably from 103.4 to 344 .73 kPa (15 to 50 psi) and a retention time of the one or more alkanes in the oxidative dehydrogenation reactor 2 riser of 0.002 to 20 seconds.
[092] The flared section 12 of the riser should be wide enough to cause the catalyst particles to fall into the fluidized bed zone 6, the spreader plate 13 should be high enough to minimize catalyst friction.
[093] Port 14 on riser 10 allows the introduction of one or more low temperature steam and atomized water (eg a mist) at a temperature of at least about 25 °C, desirably 50 °C lower at the temperature of the oxidative dehydrogenation reactor. In some embodiments, the steam has a temperature of from about 200°C to about 400°C, in additional embodiments, the temperature can be from about 300°C to 350°C. The steam cools the catalyst coming from the regenerator reactor 4 and also removes any entrained or absorbed impurities (eg ethylene or air). The atomized water may have a temperature of 50°C to 75°C upon introduction to port 14. All or part of the atomized water may be recycled water from the reaction product.
[094] Port 15 at or towards the bottom of the riser of oxidative dehydrogenation reactor 2 is an inlet for the hydrocarbon feed, typically high purity ethane mixed with oxygen or an oxygen containing gas. Hydrocarbon and oxygen feeds can be combined near and upstream of the oxidative dehydrogenation reactor. As it is a fluidized bed reactor, it is necessary that the upward flow of hydrocarbon and oxygen-containing gas feed is sufficiently well distributed to fluidize the bed of catalyst particles to minimize hot spots.
[095] The process of the present invention can be used to generate ethylene from relatively pure raw material.
[096] Ethane individually should comprise about 95% by weight of ethane, preferably 98% by weight of ethane and no more than about 5% by weight of associated hydrocarbons such as methane. Preferably, the feed is oxygen having a relatively high purity, in some embodiments above 90% purity, in additional embodiments above 95% purity. Although air can be used as an oxygen source, it can give rise to downstream separation issues.
[097] In the additional embodiment of the invention, the reactor of the present invention can be used to replace an ethane/ethylene separator or refinery tail gas or other hydrocarbon processing process, in which case the raw material may comprise from 10 to 80% by volume ethylene and ethane balance.
[098] To maintain a viable fluidized bed, the mass gas flow rate through the bed should be above the minimum flow required for fluidization, and preferably from about 1.5 to about 10 times Umf, and more preferably, from about 2 to about 6 times Umf. Umf is used in accepted form as the abbreviation for the minimum mass gas flow required to achieve fluidization, C. Y. Wen and Y. H. Yu, "Mechanics of Fluidization", Chemical Engineering Progress Symposium Series, Vol. 62, p. 100-111 (1966). Typically, the required surface gas velocity ranges from 0.3 to 5 m/s.
[099] At the upper end of the oxidative dehydrogenation reactor, below the withdrawal zone 9, is the port 16, which allows the spent catalyst stream to settle and leave the reactor. At the top of reactor 3, there are cyclones 17 to remove any fine catalyst particles that have not settled in the withdrawal zone 9.
[0100]The average retention time of the catalyst supported in the riser of oxidative dehydrogenation reactor 2 is less than about 30 seconds, in some cases less than 15 seconds, in some cases 1 to 6 seconds. Port 16 connects downpipe 18 with oxidative dehydrogenation reactor 3 and regeneration reactor 4. Port 19 on downfall tube 18 is positioned close to regeneration reactor 4. Port 19 allows the introduction of steam into a temperature from about 300°C to 500°C, in some embodiments, from 350°C to 450°C to flow countercurrent to the spent catalyst stream to remove raw material and entrained product. In some cases, steam can also burn surface coke into catalyst particles. The flow rate of steam in the downpipe must be low enough to prevent the supported catalyst from being returned to the take-off zone 9.
[0101]The regeneration reactor is also a fluidized bed reactor. Port 20 on the bottom of the regeneration reactor allows air and, in some cases, recycled cooled nitrogen to return to the reactor. The regeneration reactor is typically operated at temperatures from 250 °C to 400 °C and pressures from 3.447 to 689.47 kPa (0.5 to 100 psi), preferably from 103.4 to 344.73 kPa (15 to 50 psi). The retention time of the supported catalyst in the regeneration reactor is less than 3 minutes. Typically, the ratio of catalyst retention time in the regenerator reactor to the retention time in the oxidative dehydrogenation reactor is not less than 3.
[0102] The port 21 in the upper portion of the regenerator reactor 4, above the fluidized bed of supported catalyst particles allows the waste gas to leave the reactor. There may be cyclones, as described for oxidative dehydrogenation reactor 3, in the upper section of regenerator reactor 4 to remove any fine catalyst particles from the nitrogen product stream. Air and optionally nitrogen, which can be cooled, are passed through the regeneration reactor. Oxygen is substantially removed from the air. The tail gas will comprise from 85 to 100% nitrogen.
[0103]The above description of the circulating fluidized bed reactor has been largely schematic. There may be multiple valves, filters, etc. on the doors. Selection of appropriate valves should be well known to those skilled in the art. Likewise, there may be adequate means of compression and ventilation used to force gases through the system. The selection of suitable fans, compressors or expanders for cooling should be known to persons of ordinary skill in the art.
[0104]It is desirable to recover as much energy as possible from the oxidative dehydrogenation reaction and the regeneration reaction. The ethylene feed and co-products (eg CO2 and CO) from the oxidative reactor are fed to separate steam generators to generate steam. Some of the steam can be recycled back into the process. Steam can be injected into the riser to cool the catalyst particles. Steam can also be injected into the downpipe to burn any coke and carry away any absorbed or adsorbed feed or products.
[0105]The oxygen-containing current passing through the regenerator is substantially depleted of oxygen at the reactor output (eg the output current comprises not less than 90% nitrogen). If nitrogen is also used as a component of the feed stream, a portion of the product stream can be recycled to the input to the regeneration reactor. The product stream portion of the regeneration reactor may be subjected to one or more cooling or cooling steps to maintain heat balance in the regenerator. In some embodiments, a CO promoter can be added to the regenerator to minimize heat release to the regenerator (ie, reduce/control CO2 production).
[0106] There are a variety of patents and applications that teach about CO promoters, including the following.
[0107] U.S. 4,064,039 issued December 20, 1977 to Penick, assigned to Mobil Oil Corporation, teaches adding up to 50, typically from 0.01 to 1 ppm of Palladium and Rhenium Group Metals to promote combustion of CO.
[0108] U.S. 4,072,600, granted February 7, 1978 to Schwartz, assigned to Mobil Oil Corporation, teaches adding 1 to 50 ppm of a metal selected from the group consisting of platinum, palladium, iridium, osmium, ruthenium and rhenium and mixtures to promote the combustion of CO.
[0109] U.S. 4,093,535, issued June 6, 1978 to Schwartz, assigned to Mobil Oil Corporation, teaches adding 1 to 50 ppm of a metal selected from the group consisting of platinum, palladium, iridium, osmium, ruthenium and rhenium and mixtures to promote the combustion of CO.
[0110] U.S. 5,565,399, issued October 15, 1996 to Fraenkel et al., assigned to Engelhard Corporation, teaches supports for CO promoters comprising aluminum microspheres having been impregnated with at least 2 percent by weight of La2O3 and from 3 to 8 weight percent CeO2 which microspheres are substantially free of alpha-aluminum and having an x-ray pattern showing the presence of crystalline CeO2.
[0111] The process of the present invention must be operated to have a conversion of not less than 80% (for ethylene) and a selectivity of not less than 90%, preferably greater than 95% for ethylene. Product Stream Separation:
[0112] The stream 22 that exits the dehydrogenation reaction comprises ethylene, water (steam - steam) and a small amount of ethane, unconsumed oxygen and waste gases, typically CO and CO2. The problem of separation needs to be considered in the context of the intended use of ethylene.
[0113]There are a variety of processes that can use dilute ethylene, such as polymerization processes. However, this approach needs to be balanced with the effect of polar molecules such as CO and CO2 and oxygen on the catalyst used for polymerization. It may be preferable to separate the polar molecules prior to separating ethylene and ethane. Polar molecules can be separated by an adsorption bed, such as a zeolite bed. In the simplest modality, depending on component ratios, the bed can be regenerated and all components fed to a burner to burn off the CO. However, in a chemical complex, there are other unit operations that can use CO as a feed (various processes of carboxylic acid and anhydride (acetic acid, methacrylic acid and maleic anhydride). If there is a significant amount of CO and CO2, the components can be separated. There are a variety of methods well known in the art to separate CO2 and CO. The stream must be cooled and washed and then passed through an adsorbent such as activated carbon (to remove impurities from CO2) or a liquid amine separator or a liquid carbonate separator to absorb CO2. CO can be separated by a variety of techniques. Depending on the volume, a vacuum separation method using activated carbon as an adsorbent may be suitable, a membrane separation may be adequate and adsorption to copper ions (on a suitable support) may be adequate. Oxygen Removal - Fixed Bed:
[0114] In one embodiment, there may be two or more fixed bed reactors, having an oxidative dehydrogenation catalyst that releases or captures oxygen, which are used as scavengers to accommodate the product flow out of the circulating fluidized bed oxidative dehydrogenation reactor . Fixed bed reactors have tubes and valves so that the feed to the fluidized bed oxidative dehydrogenation reactor passes through one or more of the fixed bed reactors having a catalyst containing oxygen that is either consumed or abandoned. This is not a major problem with the prereactor operating in the oxidative dehydrogenation mode, as any excess undehydrogenated alkane in the prereactor will be converted into the fluidized bed oxidative dehydrogenation reactor. The key issue is catalyst depletion in the oxygen fixed bed reactor. The tubes and valves flow the product stream through one or more of the fixed bed reactors having oxidative dehydrogenation catalysts that are oxygen depleted. The depleted fixed bed catalyst scavenges oxygen from the product stream. As noted, the stream valves and tubes can be operated so that the feed streams flow through the oxygenated fixed bed catalyst reactor and the product stream flows through one or more of the oxygen depleted fixed bed catalyst reactors.
[0115]Oxygen-containing oxidative dehydrogenation catalyst may have oxygen as matrix oxygen, adsorbed oxygen, or adsorbed oxygen on the catalyst, the support, or both. The oxygen-depleted oxidative dehydrogenation catalyst has reduced oxygen, preferably about 60% less oxygen in the catalyst, and supported as matrix oxygen, adsorbed oxygen or adsorbed oxygen on the catalyst, on the support, or both.
[0116] Preferably, at the outlet of the fluidized bed oxidative dehydrogenation reactor is an oxygen sensor. Additionally, there must be an oxygen sensor at the outlet for the dehydrogenated product of each fixed bed reactor to determine the level of oxygen leaving the product leaving that fixed bed reactor. When the oxygen level increases in the dehydrogenated product output of the fixed bed reactor operating in sequestrant mode, this indicates that the catalyst has substantially absorbed reactive oxygen (and can be returned for use as a pre-reactor). The amount of reactive oxygen absorbed by the oxygen-depleted catalyst in pre-reactor operation in chemoabsorption or oxygen sequestrant mode should not be less than about 1.5%, typically about 2% of the total oxygen in the catalyst, (this will also correspond to the amount of reactive oxygen available for catalyst release in pre-reactors in oxidative dehydrogenation mode).
[0117] A mode for operating using three pre-reactors is schematically illustrated in Figures 10, 11 and 12 (where equal parts have equal numbers) and in the table below. In Figures 10, 11 and 12, the valves are not shown. The configuration of the main reactor is the same, however, the bypass of the valves makes the pre-reactor, hijacking reactor and protection reactor seem to “switch” places. A pre-reactor operates in this way and converts part of the feed stream to ethylene. An oxygen-depleted prereactor acts as a chemoabsorption reactor or primary oxygen sequestrant and a second prereactor (also oxygen depleted acts as a secondary or shield oxygen scavenging or chemoabsorption reactor).



[0118]In an alternative embodiment, oxygen can be separated from the product stream using cryogenic methods. However, this adds investment and operating costs to the process.
[0119] The above sequestering process is more fully described in Canadian Patent Application No. 2,833,822, filed November 21, 2013, the text of which is incorporated herein by reference.
[0120]Residue gases from the downpipe are also subject to the same separation techniques to recover them.
[0121] As noted above, it may not be necessary to separate ethane from ethylene at this stage, however, if desired, there are a variety of techniques that can be used.
[0122]The most common techniques would be to use a cryogenic C2 separator. Other separation techniques include the following.
[0123]One method of separating the product stream is by absorption. The gaseous product stream comprising mainly ethane and ethylene may be contacted in a countercurrent flow with a heavier paraffinic oil, such as mineral seal oil or medicinal white oil at a pressure of up to 800 psi (about 5.5x103 kPa) and at temperatures from about 25°F to 125°F (about -4°C to about 52°C). Ethylene and low boiling components are not absorbed into the oil. Ethane and high-boiling components are absorbed into the oil. Ethylene and low boiling components can then be passed to the C2 separator if necessary. The absorption oil can be selectively extracted with a solvent such as furfural, dimethyl formamide, sulfur dioxide, aniline, nitrobenzene and other known solvents to extract any heavier paraffins. This process is described in greater detail in United States Patent No. 2,395,362, issued May 15, 1945 to Welling, assigned to Phillips Petroleum Company, the contents of which are incorporated herein by reference.
[0124]Another separation method is an adsorption method. The adsorbent preferably adsorbs one of the components in the product stream. The adsorption method typically comprises a train of two or more adsorption units, so that when one unit has reached capacity, the feed is directed to an alternate unit, although the fully recharged unit is typically regenerated by one or more than one unit. change in temperature or pressure or both.
[0125] There is a significant amount of technique in separating ethylene and ethane using silver or copper ions in their +1 oxidation state. Olefins are preferentially absorbed in a complexing solution containing the complexing agent selected from silver salts (I) or copper (I) dissolved in a solvent. Some silver absorbents include silver nitrate, silver fluoroborate, silver fluorosilicate, silver hydroxyfluoroborate and silver trifluoroacetate. Some copper absorbents include cuprous nitrate; cuprous halides such as cuprous chloride; cuprous sulfate; cuprous sulfonate; cuprous carboxylates; cuprous salts of fluorocarboxylic acids such as cuprous trifluoroacetate and cuprous perfluoroacetate; cuprous fluorinated acetylacetonate; cuprous hexafluoroacetylacetonate; cuprous dodecylbenzenesulfonate; copper-aluminum halides such as cuprous aluminum tetrachloride; CuAlCH3Cl3; CuAlC2 H5Cl3; and cuprous aluminum cyanotrichloride. If the product stream has been dried prior to contact with the liquid adsorbent, the adsorbent must be stable to hydrolysis. The complexing agent is preferably stable and has high solvent solubility. After an adsorbent solution is substantially charged, the product current supply is switched to an additional solution. The adsorbent solution that is fully charged is then regenerated through changes in heat or pressure or both. This releases the ethylene.
[0126] These types of processes are described in U.S. Patents No. 6,581,476 issued February 11, 2003 to Culp et al., assigned to Union Carbide Chemicals & Plastics Corporation, and No. 5,859,304, issued in January 12, 1999 to Barchas et al., attributed to Stone and Webster Engineering, the contents of which are hereby incorporated by reference.
[0127] In an alternative to solution, process supports, such as zeolite 4A, zeolite X, zeolite Y, alumina and silica, can be treated with a copper salt to selectively remove carbon monoxide and/or olefins from a gaseous mixture containing saturated hydrocarbons (ie, paraffins) such as ethane and propane. United States Patent No. 4,917,711, issued April 17, 1990 to Xie et al., assigned to Peking University, describes the use of such supported adsorbents, the contents of which are incorporated herein by reference.
[0128] Likewise, United States Patents No. 6,867,166, issued March 15, 2005, and No. 6,423,881 issued July 23, 2002 to Yang et al., assigned to Regents of the University of Michigan, which are hereby incorporated by reference, describe the use of copper salts and silver compounds alternatively supported on silica, alumina, zeolite MCM-41, zeolite 4A, carbon molecular sieves, polymers such as Amberlyst-resin 35, and alumina for selectively adsorbing olefins from gas mixtures containing olefins and paraffins. Kinetic and thermodynamic separation behaviors were observed and modeled. Olefin adsorption occurs at pressures from 1 to 35 atmospheres, preferably less than 10 atmospheres, more preferably less than 2 atmospheres at temperatures from 0 to 50 °C, preferably from 25 to 50 °C and desorption occurs at pressures from 0.01 to 5 atmospheres, preferably 0.1 to 0.5 at temperatures from 70°C to 200°C, preferably from 100°C to 120°C.
[0129] In a further embodiment, the adsorbent can be a physical adsorbent selected from the group consisting of natural and synthetic zeolites without a silver or copper salt.
[0130]In general, the adsorbent can be alumina, silica, zeolites, carbon molecular sieves etc. Typical adsorbents include alumina, silica gel, carbon molecular sieves, zeolites such as type A and type X zeolite, type Y zeolite etc. The preferred adsorbents are zeolite type A and the most preferred adsorbent is zeolite type 4A.
[0131] Type 4A zeolite, ie the sodium form of type A zeolite, has an apparent pore size of about 3.6 to 4 Angstrom units. This adsorbent provides greater selectivity and adsorption capacity of ethylene from blends and ethylene-ethane and propylene from propylene-propane blends at elevated temperatures. Such adsorbent is most effective for use in the invention when it is substantially unmodified, that is, when it has only sodium ions as its exchangeable cations. However, certain properties of the adsorbent, such as thermal and light stability, can be improved by partially exchanging some of the sodium ions with other cations (other than silver and copper). Therefore, it is within the scope of the preferred embodiment of the invention to use a type 4A zeolite in which some of the sodium ions coupled to the adsorbent are replaced with other metal ions, provided that the percentage of exchanged ions is not so large that the adsorbent loses its Type 4A feature. Among the properties that define the Type 4A characteristic are the ability of the adsorbent to selectively adsorb ethylene from ethylene-ethane and propylene mixtures from propylene-propane gas mixtures at elevated temperatures, and to achieve this result without causing oligomerization or significant polymerization of the alkenes present in the mixtures. In general, it has been determined that up to about 25 percent (on an equivalent basis) of the sodium ions in zeolite 4A can be replaced by ion exchange with other cations without stripping the adsorbent of its 4A-type characteristic. Cations that can be ionically exchanged with the 4A zeolite used in the alkene-alkane separation include, among others, potassium, calcium, magnesium, strontium, zinc, cobalt, manganese, cadmium, aluminum, cerium, etc. when exchanging other cations for sodium ions, it is preferred that less than 10 percent of the sodium ions (on an equivalent basis) be replaced with such other cations. Substitution of sodium ions can modify the properties of the adsorbent. For example, replacing some sodium ions with other cations can improve the stability of the adsorbent. As disclosed in U.S. Patent No. 5,744,687, issued April 28, 1998, to Ramachandran et al., assigned to BOC Group, Inc., the contents of which are incorporated herein by reference.
[0132]A particularly preferred zeolite is ZSM-5.
[0133] In addition to zeolites, there are a variety of titanosilicate homologues referred to as ETS compounds.
[0134] United States Patent No. 5,011,591 describes the synthesis of a titanosilicate with large pore diameter designated "ETS-10". In contrast to ETS-4 and CTS-1 (referenced below), the large pore titanosilicate material, ETS-10, which has pore diameters of about 8 A, cannot kinetically distinguish light olefins from paraffins of the same number of carbon. However, high degrees of selectivity have been reported for the separation of ethylene from ethane using the prepared ETS-10 zeolites; see: Al-Baghli and Loughlin in J. Chem. Eng. Data 2006, v51, p 248. The authors demonstrate that Na-ETS-10 is capable of selectively adsorbing ethylene from a mixture of ethylene and ethane under thermodynamic conditions, even at room temperature. Although the reported selectivity for ethylene adsorption using Na-ETS-10 was high at room temperature, the adsorption isotherms for ethylene and ethane had highly rectangular shapes consistent with a low pressure swing capability. Consequently, Na-ETS-10 is not readily applicable to pressure swing absorption (PSA) processes, at least at room temperature or below.
[0135]However, the cationic modification of Na-ETS-10 as prepared provides an adsorbent for the PSA separation of olefins and paraffins having the same number of carbon atoms, at room temperatures. Mono-, di- and trivalent cations are selected from the group of compounds of 2-4 metals, a proton, ammonium and mixtures thereof. Some specific non-limiting examples of mono, di-, or trivalent cations that can be used in the present invention include, Li+, K+, Cs+, Mg2+, Ca2+, Sr2+, Ba2+, Sc3+, Y3+, La3+, Cu+, Zn2+, Cd2+, Ag+ , Au+, H+, NH4+ and NR4+, where R is an alkyl, aryl, alkylaryl or arylalkyl group. Cationic modifiers are usually added to unmodified Na-ETS-10 in the form of a salt or an acid. The anionic counterion associated with the cationic modifier is not specifically defined as long as it does not adversely affect the modification reactions (ie, cation exchange). Suitable anions include, but are not limited to, acetate, carboxylate, benzoate, bromate, chlorate, perchlorate, chloride, citrate, nitrate, nitride, sulfates and halide (F, Cl, Br, I) and mixtures thereof. Suitable acids include organic and inorganic acids, with organic acids being preferred. United States Patent 8,017,825, issued September 13, 2011 to Kuznicki et al., assigned to the Governors of the University of Alberta, describes the technology, the text of which is incorporated herein by reference.
[0136] As described in U.S. Patent No. 6,517,611, heat treatment of ETS-4 generated a zeolite material with controlled pore volume, designated "CTS-1", which is a highly selective absorbent for separations. olefin/paraffin. Zeolite CTS-1, which has pore diameters of about 3-4 A, selectively adsorbed ethylene from a mixture of ethylene and ethane through a size exclusion process. The pore diameter of CTS-1 allowed the diffusion of ethylene while blocking the diffusion of ethane which was too large to enter the pores of the CTS-1 zeolite, thus providing a kinetic separation. The CTS-1 adsorbent has been successfully applied to a PSA process, where ethylene or propylene can be separated from ethane or propane respectively.
[0137]The above adsorbents can be used in pressure swing adsorption units. Typically, the absolute pressure range used during the adsorption step can be from about 10 kPa to about 2,000 kPa, (about 1.5 to about 290 pounds per square inch (psi)), preferably about about 50 kPa to about 1000 kPa (from about 7.2 to about 145 psi). The range of pressures used during the release of adsorbate (ie during the regeneration step) can be from about 0.01 kPa to about 150 kPa (about 0.0015 to about 22 psi), preferably from about 0.1 kPa to about 50 kPa (about 0.015 to about 7.3 psi). In general, the adsorption step can be carried out from ambient temperatures to above about 200 °C, preferably below 150 °C, more preferably below 100 °C, provided the temperatures do not exceed the temperatures in that chemical reactions of the olefin occur, such as an oligomerization or polymerization.
[0138]Another class of adsorbents is ionic liquids. Olefins and paraffins can be separated using ionic liquids of the formula, a metal dithiolene selected from the group of complexes of the formulas:
where M is selected from the group consisting of Fe, Co, Ni, Cu, Pd and Pt; and R1, R2, R3, R4, R5, and R6 are independently selected from the group consisting of a hydrogen atom, electron withdrawing groups including those that are or contain heterocyclic, cyano, carboxylate, carboxylic ester, keto, nitro groups and sulfonyl, hydrocarbyl radicals selected from the group consisting of C1-6 alkyl groups, C5-8 alkyl groups, C2-8 alkenyl groups and C6-8 aryl groups whose hydrocarbyl radicals are unsubstituted or fully or partially substituted, preferably those substituted. by halogen atoms. The ionic liquid can be used with a co-solvent or non-reactive solvent. The solvent can be selected from the group of conventional aromatic solvents, typically toluene. Adsorption pressures can range from 200 psig to 300 psig (1.3x103 to 2x103 kPag), preferably below 250 psig (1.7x103kPag) and adsorption temperatures can range from ambient to 200 °C, preferably below 150 ° C, and the olefin can be released from the ionic liquid by one or more pressure reduction by at least 50 psi (3.4x102 kPa) and temperature increase by not less than 15°C.
[0139]Nitrogen from the regeneration reactor, not recycled to the regeneration reactor, can be used in a number of downstream unit operations. Potential downstream unit operations include an ammonia plant, an acrylonitrile plant, a urea plant and an ammonium nitrate plant.
[0140] The following non-limiting examples demonstrate the present invention.
[0141] The catalyst used in the experiments was the formula: MoaVbNbcTeeOd, where: a is from 0.90 to 1.10; b is from 0.25 to 0.4; c is from 0.1 to 0.3 ;e is from 0.1 to 0.3; and d is a number to satisfy the valence state of the mixed oxide catalyst.
[0142]The reactor used in the experiments consisted of a quartz tube reactor. The sample size was typically about 0.5 cm 3 , 0.17 g. The particle size for the catalyst was 0.2 - 0.7 mm.
[0143]The reactor was initially operated in a regeneration mode (oxidation of the catalyst). The reactor was heated to a temperature of 355 °C to 397 °C in air for 30 minutes. Then, the gas flow was switched to a mixture of 75% by volume of ethane and 25% by volume of oxygen. The flow rate of the mixture of ethane and oxygen varied by 300/600/1200 cm3 (Stp) per hour. The reaction took place during the first minute of the passage of the reactants over the oxidized catalyst bed. The catalyst bed was then reoxidized and then a mixture of ethane and oxygen was passed through the oxidized catalyst. The gas leaving the reactor was analyzed to measure the residual oxygen and the amount of ethane, ethylene and by-products in the product gas.Experiment No.1:(air «• gas mixture [75%C2H6+25%O2]) 355 °C
[0144] Figures 2 and 3 demonstrate the time dependence of ethane and O2 conversion, as well as the selectivity of ethylene formation when the gradual reduction of the pre-oxidized catalyst by the reaction mixture provided at 600 cc/h in two different temperatures.
[0145]It can be seen (Figures 2, 3) that all transient processes occur during the first minute of operation under our test conditions. The effect is not pronounced at ~355°C, only a slight increase in conversion without any loss of selectivity can be noticed (Figure 2). The same effect of increased conversion becomes stronger at 400 °C, but in this case, it is accompanied by a substantial loss of selectivity due to additional CO2 formation (Figure 3). It should be mentioned that the process that takes place most actively at 400 °C is accompanied by measurable self-heating of the catalyst layer (~5-6 °C measured at the reactor wall). Some unwanted complete oxidation contribution in the gas phase cannot be excluded. Experiment 2:
[0146]Experiment 1 was repeated at 398 °C.
[0147] Comparing Experiments 1 and 2, the maximum conversion was raised above 70%. Experiment 3:
[0148] To clarify the invention, an additional test was performed with varying gas flow rates (300 and 1200 cc/min) using the same condition as in examples 1 and 2, except that the flow rate was 1200 cm3 /H. The results obtained are shown in Figures 4 and 5.
[0149]The data obtained (Figures 3-5) show that the selectivity is related to the feed flow rate (spatial velocity). Reducing the gas flow rate below 600 h-1 generated a temporary drop in selectivity below 75% (Figure 5b). Again, the process is accompanied by considerable self-heating of the catalytic layer after switching from gas to the reaction mixture (~6-7 °C measured at the reactor wall). The selectivity curves are summarized and compared in Figure 6. It is interesting to note that in the short retention time, despite the high conversion, very little oxygen in the gas phase is consumed. Thus, the undesirable contribution of complete oxidation with temporary heating of the catalyst bed becomes more and more pronounced as the contact time increases (Figure 6). At the same time, increasing the gas flow rate up to 2400 h-1 allows us to avoid a considerable contribution from total oxidation (Figure 6).
[0150]For quantitative comparison of the conversion data, all results obtained at flow rates differing by a factor of 2 are presented in Figure 7 using an absolute scale (ie, as a function of the amount of ethane fed through the reactor) . All three curves are very similar (Figure 7). It is evident that the high initial conversion of ethane (70-80%) is caused by the presence of extra oxygen stored in the pre-oxidized catalyst, and the transient process shown in Figure 7 is related to the gradual loss of this additional oxygen. The results obtained allow us to calculate the amount of “reactive” matrix oxygen involved in the reaction during the transient process. Catalyst oxygen depletion is the same for all three tests and can be evaluated as ~1% of the total matrix oxygen of our Mo-V-Te-Nb-Ox catalyst. Experiment 4: Periodic redox cycle (pure O2 «• C2H6) pure): benchmark test
[0151]To clarify the invention, this experiment was carried out in the absence of oxygen in the ethane stream under the same other conditions (ie, flow rate and temperature). In this test, the catalyst charge placed in a quartz reactor was heated to a given temperature (354 °C or 397 °C) in pure O2 flow, maintained for 30 min, then gas flow (600 cm3/h) was switched to pure C2H6, and the exit gas sample was analyzed after a certain time. After reoxidation of the catalyst for 30 min, measurements were repeated several times with varying time intervals, and response curves resulting from products were received (up to 3 min). Figures 8 and 9 demonstrate the dependence of ethane conversion time and residual O2 content, as well as selectivity of ethylene formation upon gradual reduction of catalyst by ethane at two different temperatures.
[0152]Transient processes occur for 1-2 minutes under our test conditions (Figures 8, 9). The reaction is accompanied by a measurable loss of selectivity. The effect is quite pronounced even at ~350 °C (Figure 8b) and becomes stronger at 400 °C (Figure 9b). It is important to note that the reaction is accompanied by measurable self-heating of the catalyst layer (4-8°C measured on the outer wall of the reactor). The switch back to O2 flow for catalyst reoxidation is also accompanied by some catalyst heating (3-4°C). Furthermore, this heating appears to be non-uniform, but mobile through the layer during the reaction. Taking into account that this catalyst overheating is considerably stronger within the catalyst bed, the role of non-isothermal conditions provided by alternation between two pure gases can be important.
[0153]The above examples also illustrate that conversions and selectivity using a pulse mode of ODH are not as effective as the present invention. INDUSTRIAL APPLICABILITY
[0154] The invention provides an oxidative dehydrogenation process using a circulating bed reactor (similar to a fluidized bed catalyst cracker (FCC)) providing good yields of olefin product in high selectivity.

权利要求:
Claims (24)
[0001]
1. Process for the oxidative dehydrogenation of one or more alkanes selected from ethane and propane in the presence of a supported catalyst selected from: i) catalysts of the formula: VxMoyNbzTemMenOpem which Me is a metal selected from Ta, Ti, W, Hf, Zr, Sb and mixtures thereof; x is from 0.1 to 3 provided that when Me is absent, x is greater than 0.5;y is from 0.5 to 1.5;z is from 0.001 to 3;m is from 0.001 to 5;n is from 0 to 2;ep is a number to satisfy the valence state of the mixed oxide catalyst on a metal oxide support; eii) catalysts of the formula: MoaVbNbcTeeOdem which: a is from 0.75 to 1.25; b is from 0.1 to 0.5; c is from 0.1 to 0.5; and is from 0.1 to 0 .35, e1. is a number to satisfy the valence state of the mixed oxide catalyst on a metal oxide support; CHARACTERIZED by the fact that the process comprises: passing through an oxidative dehydrogenation reactor comprising a riser containing a fluidized bed of said catalyst said one or more alkanes and oxygen at a temperature of 250 °C to 450 °C, a pressure of 3.447 to 689.47 kPa (0.5 to 100 psi) and a retention time of said one or more alkanes in said riser of 0.002 to 10 seconds, and reducing said catalyst, said catalyst having an average retention time in the dehydrogenation reactor of less than 30 seconds; feeding said reduced catalyst through a said drop tube of said dehydrogenation reactor to a separate fluidized bed regeneration reactor and passing an air stream optionally with additional nitrogen at a temperature of 250 °C to 400 °C and pressures from 3.447 to 689.47 kPa (0.5 to 100 psi) across said bed to oxidize said catalyst; and passing said oxidized catalyst from said fluidized bed regenerator back to said oxidative dehydrogenation riser together with steam at a temperature of 300°C to 350°C or atomized water at a temperature of 50°C to 75°C or both steam at a temperature of 300 °C to 350 °C and water at a temperature of 50 °C to 75 °C; wherein the amount of oxygen in the feed to said riser is 18 mol% to 26% mol and the conversion of alkane to alkene is not less than 50% per step and the selectivity for the conversion of alkane to alkene is not less than 0.9.
[0002]
2. Process according to claim 1, CHARACTERIZED by the fact that the upper part of said riser comprises an inverted cone disperser.
[0003]
3. Process according to claim 2, CHARACTERIZED by the fact that it further comprises passing steam at a temperature of 350 °C to 450 °C in countercurrent to the flow of oxygen-depleted catalyst through said drop tube.
[0004]
4. Process according to claim 3, CHARACTERIZED in that it further comprises passing air or a mixture of air and nitrogen through the regeneration reactor and generating a gas product stream comprising not less than 80 - 100% by volume of nitrogen.
[0005]
5. Process according to claim 4, CHARACTERIZED by the fact that it further comprises recycling a portion of the reduced oxygen effluent stream from the regenerator reactor and optionally cooling it and recycling it to the regenerator reactor.
[0006]
6. Process according to claim 4, CHARACTERIZED by the fact that it further comprises adding a CO promoter to the regenerating reactor.
[0007]
7. Process according to claim 5, characterized in that it further comprises separating said alkene product from the oxidative dehydrogenation reactor from the water in the product stream of the oxidative dehydrogenation unit.
[0008]
8. Process according to claim 7, CHARACTERIZED by the fact that it further comprises passing unused nitrogen from the effluent stream of the catalyst regeneration reactor to an integrated site unit operation selected from an ammonia plant, an acrylonitrile plant , a urea plant and an ammonium nitrate plant.
[0009]
9. Process according to claim 1, CHARACTERIZED by the fact that the retention time of the catalyst in the oxidative dehydrogenation reactor is less than 30 seconds.
[0010]
10. Process according to claim 9, CHARACTERIZED by the fact that the retention time of the catalyst in the regeneration reactor is less than 3 minutes.
[0011]
11. Process according to claim 10, CHARACTERIZED by the fact that the ratio of the catalyst retention time in the regeneration reactor to the catalyst retention time in the oxidative dehydrogenation reactor is not less than 3.
[0012]
12. Process according to claim 11, CHARACTERIZED by the fact that the product stream from the oxidative dehydrogenation reactor and at least a portion of the effluent stream from the regenerator reactor are passed through separate steam generators for heat recovery.
[0013]
13. Process according to claim 12, CHARACTERIZED by the fact that the product stream from the oxidative dehydrogenation reactor is cooled and passed through a column to separate alkene combustion products.
[0014]
14. Process according to claim 12, CHARACTERIZED by the fact that the product stream from the oxidative dehydrogenation reactor is cooled and passed through an amine unit to remove CO2.
[0015]
15. Process according to claim 1, CHARACTERIZED by the fact that the support is selected from silicon dioxide, fused silicon dioxide, aluminum oxide, titanium dioxide, zirconium dioxide, thorium dioxide, lanthanum oxide, magnesium oxide, calcium oxide, barium oxide, tin oxide, cerium dioxide, zinc oxide, boron oxide, yttrium oxide.
[0016]
16. Process according to claim 1, CHARACTERIZED by the fact that the alkane is ethane.
[0017]
17. Process according to claim 16, CHARACTERIZED by the fact that the catalyst is of the formula: MoaVbNbcTeeOdem that: a is 0.95 to 1.1; b is 0.3 to 0.35; c is 0.1 to 0.15; and is 0.1 to 0.25; d is a number to satisfy the valence state of the mixed oxide catalyst on a metal oxide support.
[0018]
18. Process according to claim 17, CHARACTERIZED by the fact that, in the oxidative dehydrogenation reactor, the conversion to ethylene is greater than 60%.
[0019]
19. Process according to claim 18, CHARACTERIZED by the fact that, in the oxidative dehydrogenation reactor, the selectivity for ethylene is greater than 95%.
[0020]
20. Process according to claim 1, CHARACTERIZED by the fact that: the oxidative dehydrogenation reactor comprises a riser, the riser containing the fluidized bed of catalyst; the reduced catalyst is fed to the fluidized bed regeneration reactor through a decision tube of said oxidative dehydrogenation reactor; the oxidized catalyst is passed back to the riser of said oxidative dehydrogenation reactor together with one or more steam at a temperature of 300°C to 350°C and atomized water at a temperature of 50°C to 75°C; and the upper part of said oxidative dehydrogenation reactor riser comprises an inverted cone disperser.
[0021]
21. Process according to claim 20, CHARACTERIZED by the fact that: said one or more alkanes and oxygen are passed through said oxidative dehydrogenation reactor riser in an amount not less than 5% by mol; and the ratio between the catalyst retention time in the regeneration reactor and the catalyst retention time in the riser of the oxidative dehydrogenation reactor is not less than 3.
[0022]
22. Process according to claim 21, CHARACTERIZED by the fact that the oxygen in the feed stream is above 18% in mol.
[0023]
23. Process according to claim 21, CHARACTERIZED by the fact that the oxygen in the feed stream is from 5% mol to 18% mol.
[0024]
24. Process according to claim 21, CHARACTERIZED by the fact that the oxygen in the feed stream is from 18% by mol to 26% by mol.

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法律状态:
2019-09-10| B06U| Preliminary requirement: requests with searches performed by other patent offices: procedure suspended [chapter 6.21 patent gazette]|

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2020-10-06| B06A| Notification to applicant to reply to the report for non-patentability or inadequacy of the application [chapter 6.1 patent gazette]|

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2021-05-18| B09A| Decision: intention to grant [chapter 9.1 patent gazette]|

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2021-06-08| B16A| Patent or certificate of addition of invention granted|Free format text: PRAZO DE VALIDADE: 20 (VINTE) ANOS CONTADOS A PARTIR DE 08/10/2015, OBSERVADAS AS CONDICOES LEGAIS. |

优先权:

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申请号 | 申请日 | 专利标题

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CA2867731A|CA2867731A1|2014-10-15|2014-10-15|High conversion and selectivity odh process|

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CA2867731|2014-10-15|

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PCT/IB2015/057711|WO2016059518A1|2014-10-15|2015-10-08|High conversion and selectivity odh process|

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