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    • 专利摘要:
    inherently safe odh operation. in the operation of an oxidative dehydrogenation (odh) process it is desirable to remove oxygen from the product stream for several reasons, including to reduce product oxidation. this can be achieved by having several pre-reactors upstream of the main reactor having a catalyst system containing unstable oxygen. the feed passes through one or more reactors saturated with unstable oxygen. when unstable oxygen is consumed through a valve system the pre-reactor accepts product from the main reactor and complexes reactive oxygen into the product stream until the catalyst system is saturated with unstable oxygen. then the reactor becomes a pre-reactor and the other pre-reactor becomes an ejector.
    公开号:BR112016010786B1
    申请号:R112016010786-1
    申请日:2014-11-10
    公开日:2021-08-17
    发明作者:Vasily Simanzhenkov；Xiaoliang Gao；Edward Christopher Foy；Leonid KUSTOV；Aleksey Kucherov；Elena Finashina
    申请人:Nova Chemicals (International) S.A.；
    IPC主号:
    专利说明:

    FIELD OF TECHNIQUE
    [001] The present invention relates to an oxidative dehydrogenation reactor and a process having improved safety. There is a significant amount of technique related to the oxidative dehydrogenation of alkanes to alkenes. Typically the process involves passing a stream of one or more alkanes over an oxidative dehydrogenation catalyst at temperatures of about 300°C to 750°C in the presence of oxygen or an oxygen-containing gas. Great care must be taken to prevent the reactor feed mixture from reaching an explosive limit. In addition, it is highly desirable to remove any residual oxygen from the product stream as this could lead to a process fire. TECHNICAL BACKGROUND
    [002] There are several US patents assigned to Petro-Tex Chemical Corporation issued in the late 1960s that reveal the use of various ferrites in a steam cracker to produce olefins from paraffins. Patents include US Patents 3,420,911 and US 3,420,912 in the names of Woskow et al. Patents teach to introduce ferrites such as zinc, cadmium, and manganese (i.e., oxides mixed with iron oxide). Ferrites are introduced into a dehydrogenation zone at a temperature of about 250°C to about 750°C at pressures less than 100 psi (689.476 kPa) for a time of less than 2 seconds, typically 0.005 to 0. 9 seconds. The reaction appears to take place in the presence of steam which can tend to shift the balance in the “wrong” direction. Furthermore, the reaction does not take place in the presence of a catalyst.
    [003] Document GB 1,213,181, which appears to correspond in part to the Petro-Tex patents above, reveals that nickel ferrite can be used in the oxidative dehydrogenation process. The reaction conditions are comparable to those of the Petro-Tex patents noted above.
    [004] In Petro-Tex patents the metal ferrite (eg, M FeO4 where, for example, M is Mg, Mn, Co, Ni, Zn or Cd) is circulated through the dehydrogenation zone and then to a zone of regeneration where the ferrite is reoxidized and then fed back to the dehydrogenation zone.
    [005] It is interesting to note that reversible ferrite absorbs and releases oxygen.
    [006] US Patent 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 of 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 the oxidative dehydrogenation of alkane (ethane).
    [007] US Patent 6,521,808 issued February 18, 2003 to Ozkan, et al., assigned to Ohio State University teaches sol-gel supported catalysts for 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.
    [008] US Patent 4,450,313, issued May 22, 1984 to Eastman et al., assigned to Phillips Petroleum Company discloses a catalyst of LiO-TiO2 composition, which is characterized by a low ethane conversion that does not exceed 10 %, despite a reasonably high selectivity for ethylene (92%). The biggest deficiency of this catalyst is the high temperature of the oxidative dehydrogenation process, which is close to or higher than 650°C.
    [009] The preparation of a supported catalyst that can be used for the oxydehydrogenation of ethane to ethylene at a low temperature is disclosed in US Patent 4,596,787 A, June 24, 1986 assigned to Union Carbide Corp. A supported catalyst for the low temperature gas phase of ethane to ethylene oxydehydrogenation is prepared by (a) preparing a precursor solution having soluble and insoluble portions of metal compounds; (b) separating the soluble portion; (c) impregnating a catalyst support with the soluble portion and (d) activating 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, T1, Pb, As, Bi, Te, U, Mn and/or W; a is 0.5 to 0.9, b is 0.1 to 0.4, c is 0.001 to 0.2, d is 0.001 to 0.1, and is 0.001 to 0.1 when X is present.
    [010] Another example of the low temperature oxydehydrogenation of ethane to ethylene using a calcined oxide catalyst containing molybdenum, vanadium, niobium and antimony is described in US Patent 4,524,236 A, June 18, 1985 and US 4,250,346 A, of February 10, 1981, both assigned to Union Carbide 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 at 220 to 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% ethane conversion.
    [011] US Patents 6,624,116 issued September 23, 2003 to Bharadwaj, et al. and US 6,566,573 issued May 20, 2003 to Bharadwaj, et al., both assigned to Dow Global Technologies Inc. disclose systems of Pt-Sn-Sb-Cu-Ag monolith that were tested in an autothermal regime at T>750°C, the starting gas mixture contained hydrogen (H2:O2 = 2:1, GHSV = 80,000 h-1). The catalyst composition is different from that of the present invention and the present invention does not consider the use of molecular hydrogen in the feed.
    [012] US Patents 4,524,236, issued June 18, 1985 to McCain, assigned to Union Carbide Corporation and US 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 a temperature of 375 to 400o C the conversion of ethane reached 70% with a selectivity close to 71 to 73%. However, these parameters were only achieved at very low gas hourly space velocities of less than 900 h-1 (ie, 720 h-1).
    [013] Patent JP 07053414 teaches a catalyst supported on silica of the formula Mo1.V0.3Nb0.12Te0.23On where n satisfies the valence of the catalyst for the dehydrogenation of ethane.
    [014] US Patent 7,319,179 issued January 15, 2008 to Lopez-Nieto et al., assigned to the Consejo Superior de Investigaciones Cientificas and Universidad Politecnica de Valencia, discloses Mo-V-Te-Nb-oxide catalysts. This provided 50 to 70% ethane conversion and up to 95% selectivity to ethylene (at 38%) conversion at 360 to 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 silica on 20 to 70% by weight of the total supported catalyst or ( ii) silicon carbide. The supported catalyst is prepared by conventional methods of precipitation from solutions, drying the precipitate and then calcining.
    [015] The preparation of a Mo-Te-V-Nb composition is described in WO 2005058498 A1, published June 30, 2005 (corresponding to published application US 2007149390A1). Catalyst preparation involves preparing a slurry by combining an inert ceramic carrier with at least one solution comprising ionic species of Mo, V, Te, and Nb, drying the slurry to obtain a particulate product, precalcining the dry product at 150-350°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 that of the unsupported catalyst.
    [016] 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 an ethylene effluent is disclosed in WO 2006130288 A1, December 7, 2006, assigned to Celanese Int. Corp. The catalyst has a selectivity for ethylene of 50 to 80% thus allowing 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 oxide, magnesium oxide, calcium oxide, barium oxide, tin oxide, cerium dioxide, zinc oxide, boron oxide, boron nitride, boron carbide, boron phosphate, zirconium phosphate, silicate of aluminum, silicon nitride, silicon carbide, and glass, carbon, carbon fiber, activated carbon, metal or metal oxide networks 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 of 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).
    [017] US Patent 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 in a support 1.5 to 2 mm thick.
    [018] US Patent 6,818,189 issued November 16, 2004 to Adris et al., assigned to SABIC, teaches in the passage connecting pages 9 and 10 a process in which ceramic pellets are wrapped around a reactor. tubular tube and different reagents flow around the outside and inside of the tube. The patent is directed to the oxidative dehydrogenation of ethane to ethylene.
    [019] US Patent 3,904,703 issued September 9, 1975 to Lo et al., assigned to El Paso Products Company teaches an oxidative reactor divided into zones or layers in which following a zone for oxidative dehydrogenation there is an “oxidation zone” followed by a dehydrogenation zone to oxidize hydrogen to water. Following the oxidation zone is an adsorption bed to remove water from the reactants before they enter a subsequent dehydrogenation zone. This is to reduce the impact of water on downstream dehydrogenation catalysts.
    [020] US Patent Application 2010/0256432 published October 7, 2010 in the name of Arnold et al., assigned to Lummus discloses in paragraphs 86 to 94 methods for removing residual oxygen from the product stream. A fuel such as hydrogen or a hydrocarbon can be added to the product stream to eliminate residual oxygen. The patent refers to a catalyst, but does not disclose its composition. As noted above it may then be necessary to treat the product stream to eliminate water.
    [021] US Patent 8,394,345 issued March 12, 2013 to Dieterile et al. assigned to BASF may be representative of the current approach. The patent is for a process for the catalytic partial oxidation of a hydrocarbon into a final product such as propene to acrolein or acrylic acid. The process appears to operate outside the explosive limits of propylene (Columns 32 - 33). In addition, the final product stream contains small amounts of oxygen (1.5 to 3.5% by volume, Col. 37 line 10). The residual oxygen content in the final product (acrylic acid) does not appear to be a concern for inventors. The patent teaches different from the present invention since instead of regenerating the catalyst in situ, it is replaced by a new catalyst.
    [022] US Patent 7,998,438 issued August 16, 2011 to Weis, assigned to GRT, Inc. teaches a coupling process for lower hydrocarbons to produce higher hydrocarbons involving halogenation (bromination) followed by a removal oxidation of halogen and coupling of the intermediate compounds to produce the final product. The patent is of interest as it teaches an advancing, retreating feed to burn off coke from a catalyst used in the process. The patent teaches differently by recycling product through a reaction zone to eliminate residual oxygen.
    [023] It has been known since at least October 7, 2010, the publication date for US Patent Application 2010/0256432 to remove residual oxygen from the product stream of an oxidative dehydrogenation process. However, the approach has been to consume oxygen by burning hydrocarbons or hydrogen. This is expensive and reduces the yields of the desired hydrocarbon and the selectivity for it.
    [024] The present invention seeks to provide a simple way to reduce the oxygen content in the product stream from an oxidative dehydrogenation reaction by passing the stream over a catalyst bed to draw oxygen from the product stream and at least partially provide a source of oxygen for the catalyst. DISCLOSURE OF THE INVENTION
    [025] The present invention provides a process for the catalytic oxidative dehydrogenation of one or more C2-4 alkanes comprising n pre-reactors for the oxidative dehydrogenation of said alkanes in the presence of a metal oxide oxidative dehydrogenation catalyst system mixture that absorbs oxygen in the catalyst, where n is an integer of 2 or more, and one or more downstream main oxidative reactors comprising: i) passing a feed stream comprising said one or more C2-4 alkanes through one or more of n-1 of the prereactors at a temperature of 300°C to 500°C and a pressure of 3.447 kPag to 689.47 kPag (0.5 to 100 psig) to oxidatively dehydrogenate at least a portion from the feed stream until the oxidative dehydrogenation catalyst is depleted of reactive oxygen; ii) bypass the feed stream from the pre-reactor(s) in which the oxidative dehydrogenation catalyst is depleted of reactive oxygen to a pre- reactor in which the oxidative dehydrogenation catalyst is substantially saturated with reactive oxygen; iii) passing the product stream from said n-1 pre-reactor(s) together with additional oxygen feed to one or more reactors downstream the a temperature of 300°C to 500°C and a pressure of 3,447 kPag to 689.47 kPag (0.5 to 100 psig) for the oxidative dehydrogenation of said one or more C2-4 alkanes; iv) removing a product stream of said one or more downstream reactors comprising corresponding C2-4 alkenes, unreacted C2-4 alkanes, unreacted oxygen and water [steam] and passing the same through one or more depleted pre-reactors of reactive oxygen to a temperature of 50°C to 300°C and a pressure of 3,447 kPag to 689.46 kPag to complex the oxygen in the product stream and increase the reactive oxygen saturation of the oxidative dehydrogenation catalyst and recover a substantially free product stream. oxygen; v) continue step iv) until a) exists or a pre-reactor more depleted of reactive oxygen than the one through which the product stream is being passed; or b) the oxidative dehydrogenation catalyst in the pre-reactor is substantially complexed with reactive oxygen; vi) shifting the product stream flow from the initially depleted reactive oxygen pre-reactor to a more reactive oxygen depleted pre-reactor; evii) if necessary, completely saturate the oxidative dehydrogenation catalyst in the pre-reactor initially depleted of reactive oxygen with reactive oxygen; and viii) align the initially oxygen-depleted pre-reactor. In a further embodiment the oxidative dehydrogenation catalyst in any reactor is independently selected from the group consisting of: 1) catalysts of the formula VxMoyNbzTemMenOpem that 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 and p is a number to satisfy the valence state of the mixed oxide catalyst11) catalysts of the 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 0.3; c is from 0.1 to 0.5, preferably from 0.1 to 0.3; and is from 0.1 to 0.3 preferably from 0.1 to 0.2, d is determined by the oxidation states of the other elements.
    [026] In an additional modality, said one or more downstream reactors are operated at a gas hourly space velocity (GHSV) that will be from 500 to 30,000 h-1, preferably greater than 1,000 h-1.
    [027] In an additional modality the prereactors are fixed bed reactors and the oxidative dehydrogenation catalyst is supported on an inert metal oxide support.
    [028] In a further embodiment said one or more downstream reactors are selected from the group consisting of fixed bed reactors, fluidized bed reactors or boiling and ceramic membrane reactors.
    [029] In an additional modality the process has a selectivity for said one or more C2-4 alkenes of more than 85%, preferably more than 95%.
    [030] In a further embodiment said one or more C2-4 alkanes is ethane. BRIEF DESCRIPTION OF THE DRAWINGS
    [031] Figure 1 is a schematic diagram of an apparatus and process flow to perform the invention.
    [032] Figure 2 is a schematic diagram of a main reactor in which oxidative dehydrogenation takes place on the surface of ceramic tubes.
    [033] Figure 3 is a schematic diagram of a main reactor in which oxidative dehydrogenation takes place on the surface of ceramic tubes.
    [034] Figure 4 shows the profile (Dynamics) of ethylene formation reaction as a function of time at 375°C and 400°C after the gas flow exchanges [air ^ ethane] by a catalyst Mo1-V0.3- Nb0.2-Te0.1-OX with 80% TiO2 as support.
    [035]Figure 5 shows the profile (Dynamics) of CO2 formation reaction as a function of time at 375°C and 400°C after the gas flow exchanges [air ethane]; by a Mo1-V0.3-Nb0.2-Te0.1-OX catalyst with 80% TiO2 as a support.
    [036] Figure 6 shows the selectivity of ethylene formation of time at 375°C and 400°C after the gas flow exchanges [air ^ ethane]; by a MOI-V0.3-Nb0.2-Te0.1-OX catalyst with 80% TiO2 as a support.
    [037] Figure 7 shows the profile (Dynamics) of the O2 removal reaction of the model gas mixture by the pre-reduced catalyst of Mo1-V0.3-Nb0.2-Te0.1-OX supported on 80% of TiO2 at 270°C and 400°C.
    [038] Figure 8 shows the profile (Dynamics) of the formation reaction of CO2 (A) and CO (B) after feeding the model gas mixture by the pre-reduced catalyst of Mo1-V0.3-Nb0.2- Te0.1-OX supported on 80% TiO2 at 270°C and 400°C.
    [039] Figures 9, 10, and 11 illustrate how a series of three fixed bed catalysts can be used to expel oxygen from the product stream in an oxidative dehydrogenation reactor. BEST MODE FOR CARRYING OUT THE INVENTION Numerical Ranges[1] Except in the operating examples or where otherwise indicated, all numbers or expressions referring to amounts of ingredients, reaction conditions, etc. used in the descriptive report and claims are to be understood as modified in all instances by the term “about”. Consequently, unless otherwise indicated, the numerical parameters shown in the specification and claims attached below are approximations which may vary depending on the properties which the present invention seeks to achieve. The minimum that is required, and not as an attempt to limit the application of the doctrine of equivalents to the scope of the claims, is that each numerical parameter be at least interpreted in light of the number of significant figures reported and by applying common rounding techniques. [two] Notwithstanding the numerical ranges and parameters that exhibit the broad scope of the invention are approximations, the numerical values shown 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.[3] In addition, it should be understood that any numerical range reported in this document is intended to include all sub-ranges included therein. For example, the range “1 to 10” is intended to include all sub-ranges between and including the stated minimum value of 1 and the stated maximum 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. Since the numerical ranges revealed are continuous, they include every value between the minimum and maximum values. Unless expressly stated otherwise, the various numerical ranges specified in this application are approximations.[4] All composition ranges expressed herein are limited in total to and do not exceed 100 percent (percent by volume or percent by weight) in practice. Where 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 readily understand, the amounts of the components actually used will conform. to a maximum of 100 percent.
    [040]In the disclosure "reactive oxygen" means oxygen absorbed by the oxidative dehydrogenation catalyst that is available to be used in the oxidative dehydrogenation reaction and removed from the catalyst.
    [041] In the disclosure the term "depleted reactive oxygen" when referring to the catalyst in the pre-reactor is not intended to mean absolute depletion of oxygen. Rather it means that the residual reactive oxygen levels in the catalyst are low enough that there is less than 25%, preferably less than 15%, more preferably less than 10% of the maximum amount of oxygen that has been absorbed by the catalyst. After giving off reactive oxygen the catalyst comprises metal oxides that do not give off oxygen.
    [042] Substantially saturated with reactive oxygen means that not less than 60%, preferably more than 70%, more preferably more than 85% of the reactive oxygen has been complexed with the oxidative dehydrogenation catalyst. The Catalyst System
    [043] There are several catalysts that can be used in accordance with the present invention. The following catalyst systems can be used individually or in combination. A person of ordinary skill in the art would understand that combinations should be tested on a laboratory scale to determine if there are any antagonistic effects when catalyst combinations are used.
    [044] A particularly useful family of catalysts comprises one or more catalysts selected from the group consisting of a mixed oxide catalyst of the formula i)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 provided that when Me is present 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 a 4.n is from 0 to 2, preferably n is 0, however when Me is present n is preferably from 0.5; a 1.5 ep is a number to satisfy the valence state of the mixed oxide catalyst; and ii) catalysts of the 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 0.3; c is from 0.1 to 0.5, preferably from 0.1 to 0.3; and is from 0.1 to 0.3, preferably from 0.1 to 0.2; and d is determined by the oxidation states of the other elements.
    [045] In a further modality in catalysts of group i) the x:m ratio is from 0.3 to 10, more preferably from 0.5 to 8, desirably from 0.5 to 6.
    [046] Generally, a solution is prepared from compounds of the metals selected for the catalyst, and either a particulate catalyst is formed or a supported catalyst is formed.
    [047] Methods of preparing catalysts are known to those skilled in the art. For example, the catalyst can be prepared by mixing aqueous solutions of soluble metal compounds such as hydroxides, sulfates, nitrates, halides, lower mono- or di-carboxylic acids (C1-5) and ammonium salts or the acid of metal per se. For example, the catalyst could be prepared by mixing solutions such as ammonium metavanadate, niobium oxalate, ammonium molybdate, telluric acid, etc. The resulting solution is then typically air dried at 100-150°C and calcined in an inert gas stream 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 (one) to 20, typically 5 to 15, usually about 10 hours. The resulting oxide is a brittle solid typically insoluble in water.
    [048]There are several ways in which the oxidative dehydrogenation catalyst can be supported.
    [049] In one embodiment, the support may have a low surface area, preferably less than 20 m2/g, more preferably less than 15 m2/g, desirably less than 3.0 m2/g for the dehydrogenation catalyst oxidative in the main reactor. For the oxygen-expelling catalyst a higher surface area is preferred, typically greater than 100m2/g. The support can be prepared by compression molding. At higher pressures, the interstices within the ceramic precursor being compressed fall apart. Depending on the pressure exerted on the support precursor, the surface area of the support may be from about 15 to 0.5 m2/g, preferably 10 to 0.5 m2/g, desirably from 5 to 0.5 m2/g plus preferably from 3.0 to 0.5 m 2 /g.
    [050] There is a safety advantage in using low surface area supports because there is a reduced probability that an interstitial space can be filled only with oxidizer providing an ignition source.
    [051]The low surface area support could be of any conventional shape such as spheres, rings, saddles, etc. These types of supports would be used in more conventional reactors where a mixed stream of gaseous reactants passes over the supported catalyst and the ethane is converted to ethylene. According to the present invention the catalysts in the pre-reactor are at least partially regenerated by passing the product stream from the main oxidative dehydrogenation reactor over them and removing residual oxygen from the product stream. If necessary, the pre-reactor could be isolated and further treated with oxygen or an oxygen-containing gas.
    [052] In an alternative embodiment described below, the catalyst in the one or more downstream reactors (main reactors) may be supported on a surface of a permeable ceramic membrane that defines at least part of the flow path for one reactant and the other Reagent flows over the opposite surface of the ceramic to allow the oxidizer and ethane to react on the ceramic surface.
    [053]It is important that the substrate is dried before use. Generally, the support can be heated to 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 mmol/g of support, preferably 0.5 to 3 mmol/g of support.
    [054]The amount of hydroxyl groups on silica can be determined according to the method disclosed by J.B. Peri and A.L. Hensley, Jr., in J. Phys. Chem., 72 (8), 2926, 1968, incorporated herein in its entirety by way of reference.
    [055]The dry support can then be compressed to the required shape by compression molding. Depending on the particle size of the support, it can be combined with an inert binder to hold the shape of the compressed part.
    [056] The support for the catalyst can be a ceramic or ceramic precursor formed from oxides, dioxides, nitrides, carbides and phosphates selected from the group consisting of silicon dioxide, fused silicon dioxide, aluminum oxide, dioxide 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 , boron phosphate, zirconium phosphate, yttrium oxide, aluminum silicate, silicon nitride, silicon carbide and mixtures thereof. If necessary, the support could include a binder to help shape it.
    [057] Preferred components for forming ceramic membranes include oxides of titanium, zirconium, aluminum, magnesium, silicon and mixtures thereof.
    [058] As noted above, the support in the main reactor should have a low surface area, preferably less than 10 m2/g, more preferably less than 5 m2/g, most preferably less than 3 m2/g. The support can be prepared by compression molding. At higher pressures, the interstices within the ceramic precursor being compressed fall apart. Depending on the pressure exerted on the support precursor, the surface area of the support should be less than 15 m2/g. The support will be porous and will have a pore volume from about 0.1 to 3.0 ml/g, typically from 0.3 to 1.0 ml/g. The pore size of ceramics can be small. The preferred pore size (diameter) ranges from about 3 to 10 nm. The small pore diameter is useful in ceramic membrane application because it helps to maintain the pressure drop across the membrane so that a break in the membrane is readily detected by a sudden change in pressure. Furthermore, the small pore diameter promotes a more even distribution of the reaction over the entire catalyzed surface of the membrane. That is, if larger pores are used, more of the oxygen tends to diffuse through the portion of the ceramic with which the oxygen-containing gas initially comes in contact. The remaining portion of the pottery is mostly unused.
    [059]The ceramic support can be prepared from the ceramic material using conventional techniques. For example, the starting material can be cleaned, washed and dried (or spray dried) or produced from a ceramic sol/gel and where necessary crushed or ground to the appropriate particle size. The powder can be subjected to beneficiation such as washing with acid or base to change the pore size of the ceramic.
    [060] The resulting powder is dried or calcined to remove associated water as noted above (hydration water, etc.) and can be formed as a suitable substrate, preferably tubular by, for example, compression molding or isostatic compaction to pressures of about 5 to 200 MPa (725 to 29,000 psi), with or without a binder, and sintering at temperatures that melt the particles (for example, at temperatures of about 0.5 to 0.75 of the melting temperature of ceramic material).
    [061]Other techniques may be used, such as tape molding or slurry gluing and subsequent “punching” to the required shape, such as circular, square or annular, etc. For example, annular sections could be “stacked” to produce a “tube”.
    [062]Although a tube is generally considered to be cylindrical, it could have any cross-sectional shapes, such as square, rectangular, hexagonal or star-shaped, etc. In the case of a non-cylindrical tube, wall sections could be made by bonding and then hermetically bonded together with the wall sections to form a central passage defined by an outer ceramic wall. Joints need to be hermetically sealed to prevent oxygen from coming into contact with the ethane feed and forming an explosive mixture. Glass cement or a ceramic cement or strip would be used for this purpose. A hermetic seal also needs to be at the ends of the tube where it enters and exits the reactor or joins the steel parts of the reactor.
    [063] In some modalities, once the ceramic tube is prepared, the catalyst can be deposited on the surface of the tube that will be in contact with the ethane.
    [064] The ceramic membrane can have a thickness of about 0.1 to 10 cm, typically 1 to 8 cm, preferably 2 to 7 cm.
    [065]Although ceramics are strong, they can break. It is preferred to have a support structure on at least one side, preferably the outside of the ceramic tube. More preferably, there is a support structure on the outside and inside of the tube. The structure should be in the form of a mesh or a web having holes therethrough to allow the oxygen-containing gas to pass through the support and the ceramic to react on the surface of the tube carrying the catalyst. The support can be any material suitable for use at reactor operating temperatures. From a cost standpoint, a steel mesh is probably more cost-effective. Preferably, the steel is stainless steel. The support structure must provide sufficient integrity to the tube to allow a reactor shutdown if the ceramic has a breach (eg, cracking, etc.).
    [066]One or more tubes are then placed inside the reactor. In one embodiment, the reactor is designed to have a plug flow of raw material (eg, primarily ethane) through a passage between the reactor jacket and the ceramic tube and a flow of oxygen-containing gas through the tube. of pottery. There are several provisions that come to mind. The reactor could comprise several shorter tubes placed contiguously to provide a tube of appropriate length. Or the design could be similar to a core jacket heat exchanger with several parallel tubes through which oxygen-containing gas is passed and an enclosed jacket providing a passage between the outer wall of the reactor and the ceramic tubes defining a path of flux into the ethane. The flow paths could be reversed (ethane inside and oxygen outside the tube).
    [067] In an embodiment of the invention the catalyst in the main reactor is in a ceramic membrane and in the case of the pre-reactor, in a high surface particulate support as described below.
    [068] An embodiment of the invention in which the catalyst in the main reactor is in the form of granular beds having a low surface area will be described in association with Figure 1. The raw material of alkane, preferably ethane, flows through a line 1 to a valve assembly 2 and through line 3 to pre-reactor 4, one of a pair of pre-reactors 4 and 17. In pre-reactors 4 and 17 there is a single fixed bed of catalyst, not shown. The bed is held in place between two porous membranes or open metal meshes of a mesh size small enough that particles do not pass out of the bed, again not shown. The feed passes through pre-reactor 4 and is oxidatively partially dehydrogenated and the catalyst bed is depleted of reactive oxygen. Partially dehydrogenated feed passes through outlet line 5 from prereactor 4 to another valve assembly 6. Partially dehydrogenated feed flows from valve assembly 6 through line 7 to the top of downstream reactor 8. oxidant, typically air or oxygen or a mixture of oxygen and an inert gas such as nitrogen or argon flows through line 9 and enters feed line 7 near the top of reactor 8. Mixed oxygen feed and partially dehydrogenated feed flows through three fixed catalyst beds 10, 11 and 12. There is a space between the catalyst beds and additional oxygen is fed by lines 11 and 13 into the space between the catalyst beds. The substantially dehydrogenated product stream containing small amounts of oxygen (typically less than 5% by volume, preferably less than 3% by volume) is fed through line 15 to a valve assembly 6. The dehydrogenated feed passes through valve assembly 6 via line 16 to pre-reactor 17 which is depleted or substantially depleted of reactive oxygen. As the product stream passes through prereactor 17, oxygen is drawn from it and the oxidative dehydrogenation catalyst becomes more saturated with reactive oxygen. The substantially oxygen-depleted product stream is fed from line 18 to valve assembly 2. Product passes from valve assembly 2 to line 19 for recovery and further processing.
    [069] When pre-reactor 4 is depleted of reactive oxygen then valve assemblies 2 and 6 are switched so that the alkane feed is fed to pre-reactor 17 and the product from reactor 8 is fed to the pre. - reactor 4 so that it becomes more saturated (charged) with reactive oxygen
    [070] In some instances (eg at startup) oxygen may be fed to the pre-reactor containing the supported catalyst on a support having a high surface area typically greater than 100m2/g, preferably greater than 150m2/g for " charge” the same with oxygen. This is more to balance reaction times between multiple prereactors so that a prereactor dehydrogenating feedstock will run long enough to allow full “loading” of a reactive oxygen depleted prereactor.
    [071]It is important to minimize the potential to oxidize the line 15 product stream and produce one or more of carbon monoxide and carbon dioxide. An oxidation such as this consumes valuable feed and product inventories, introduces unwanted by-products, and reduces conversion and process selectivity. To minimize additional unwanted oxidation of feed and product it is important that temperatures in the pre-reactor when adsorbing oxygen from the product stream (eg, chemisorption) are kept below the temperature for oxidative dehydrogenation (eg, from 50°C to about 300°C, preferably less than 270°C). In the pre-reactor during the chemisorption or expulsion process of oxygen from the product stream the temperature should be below about 270°C, preferably from 50°C to about 270°C, typically from 100°C to 250 °C. In view of the temperature difference between the prereactors in the oxidative dehydrogenation mode and the chemisorption or oxygen expulsion mode it may be necessary to cool the feed to the prereactor to be used for chemisorption or oxygen removal to an appropriate temperature before entering the prereactor in chemisorption or oxygen scavenging mode. It is therefore preferred to have several pre-reactors to allow the pre-reactor time to cool down before putting it into service for removal. Oxygen removal is exothermic and the reactor will heat up and depending on the catalyst system, the oxygen release could be exothermic so the thermal requirements may not be very significant (eg almost nil).
    [072] Figure 2 shows an embodiment of a main reactor comprising a membrane oxidative dehydrogenation reactor (ceramic tube). The reactor is shown generally as 30. The reactor comprises an inlet 31 into which a stream of ethane or an ethane-containing gas 32 flows. The ethane passes through the ceramic membrane tubes 33 to a collector 34. Oxygen or an oxygen containing gas 35 is fed into the tube bundle so that oxygen is outside the tubes. The ethane or ethane-containing gas 32 reacts with oxygen as it passes down the tube to form ethylene. Ethylene is collected in the collector (vertically) 34 and exits the reactor at 36.
    [073] Figure 3 shows an additional embodiment of a main reactor comprising a ceramic membrane in which ethane or ethane-containing gas 40 enters the reactor generally shown at 41 through an inlet or 42. Oxygen or o oxygen-containing gas 43 enters a tube and liner type plate shown at 44. There is a series of ceramic membrane tubes 45 encased in a steel liner 46. Ceramic membrane tubes 45 extend to the top 42 As a result, ethane or ethane-containing gas 40 flows through the interior of the ceramic membrane tubes and oxygen flows through the annular space between the exterior of the ceramic membrane tube 45 and the steel liner 46. The ethane is converted to ethylene and exits the ceramic membrane tubes into the collector (upright) 47 and exits at 48. An advantage of this design is that if a ceramic membrane loses integrity, excess oxygen enters only that tube. This is easily detected by an oxygen detector (not shown) which can be at the outlet of each tube 45 or in the manifold 47. Then the reactor can be safely shut down and the damaged tube can be located.
    [074]Reactant flows can be upstream or upstream (eg, ethane to the outside of the tube and oxygen to the inside of the tube).
    [075]The feed to the reactor comprises two separate flows to opposite sides of a tube. In a further modality, a flow, preferably to the inner surface of the tube, is an oxygen-containing gas that is selected from the group consisting of oxygen, mixtures comprising from 100 to 21% by volume of oxygen and from 0 to 79% by volume of one or more inert gases. Some inert gases can be selected from the group consisting of nitrogen, helium and argon and mixtures thereof. The oxygen containing gas could be air. The reaction
    [076] Oxidative dehydrogenation in the main reactor can be conducted at temperatures from 300°C to 550°C, typically from 300°C to 500°C, preferably from 350°C to 450°C, at pressures of 0. 5 to 100 psi (3.447 to 689.47 kPa), preferably 15 to 50 psi (103.4 to 344.73 kPa), and the residence time of paraffin (eg, ethane) in the reactor is typically from 0.002 to 30 seconds, preferably from 1 to 10 seconds. The paraffin feed (eg ethane) should preferably be 95% pure, more preferably 98% pure. Preferably, the process has a selectivity for olefin (ethylene) greater than 95%, preferably greater than 98%. The hourly space velocity of the gas (GHSV) will be from 500 to 30,000 h-1, preferably greater than 1,000 h-1. The space-time ethylene yield (productivity) in g/hour per kg of catalyst should not be less than 900, preferably greater than 1,500, more preferably greater than 3,000, most desirably greater than 3,500 at 350 to 400° Ç. It should be noted that catalyst productivity will increase with increasing temperature until selectivity is sacrificed.
    [077] The conversion of ethane to ethylene should not be less than 80%, preferably greater than 90%, most preferably 95% or greater.
    [078]The oxygen supply can be pure oxygen, however this is expensive. The feed may comprise about 95% by volume of oxygen and about 5% by volume of argon. This stream is a by-product of nitrogen production and is relatively inexpensive. Argon, being inert, should not interfere with any downstream reactions. Oxygen Expulsion
    [079]The amount of oxygen that is entrained in the ethylene product stream must be minimized for further processing. However, there will likely be some small amount of oxygen in the product stream. It is highly desirable that oxygen be removed from the product stream prior to further processing of the product stream. Immediately downstream of the oxidative dehydrogenation reactor there may be a low temperature pre-reactor (below about 270°C) in which the oxidative dehydrogenation catalyst has a reduced reactive oxygen content to absorb residual oxygen from the product stream without oxidizing more than about 5% by weight, preferably less than 1 percent by weight of the ethylene produced. The low temperature oxygen scavenging reactor operates at temperatures less than or equal to 300°C, typically 50°C to 300°C, more preferably 50°C to 270°C, generally 50°C to 270°C , preferably from 50°C to 250°C, desirably from 100°C to 250°C.
    [080]In operation it will be necessary to balance the oxygen supply to the main reactor depending on the conversion in pre-reactor 4
    [081]There may be several “pre-reactors” also used as slugs to accommodate the flow of product outside the main reactor. It may not be as problematic for the pre-reactor to operate in oxidative dehydrogenation mode as any excess undehydrogenated alkane in the pre-reactor will be converted into the main reactor(s). The key issue is the expulsion of oxygen from the product stream.
    [082] Preferably, at the output of the main oxidative dehydrogenation reactor is an oxygen sensor. In addition, there must be an oxygen sensor at the outlet for the dehydrogenated product of each pre-reactor to determine the level of oxygen leaving the process chain. When the oxygen level rises in the dehydrogenated product output of the pre-reactor (ie, knock-out reactor) this indicates that the catalyst has substantially absorbed the reactive oxygen (and can be returned for use as a pre-reactor). The amount of reactive oxygen uptake by the oxygen-depleted catalyst in the operation of the pre-reactor in oxygen-scavenging or chemisorption 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 release from the catalyst in pre-reactors in oxidative dehydrogenation mode).
    [083]A mode for operation using three prereactors is illustrated schematically in Figures 9, 10, and 11 (in which similar parts have similar numbers) and in the table below. In Figures 9, 10, and 11 the valves are not shown. The configuration of the main reactor is the same, however, changing the valves makes the pre-reactor, kick-off reactor and after reactor appear to “switch” places. A pre-reactor operates like this and converts part of the feed stream into ethylene. An oxygen-depleted pre-reactor acts as an oxygen scavenger or primary chemisorption reactor and a second pre-reactor (also depleted of oxygen acts as a secondary or posterior oxygen scavenger or chemisorption reactor).


    [084]This mode of operation is beneficial because the efficiency of most adsorption/chemisorption processes is limited by the front end or the mass transfer zone (MTZ). As a result, a significant part of the oxygen scavenging material remains unsaturated with oxygen, and consequently once the expellers are switched to pre-reactor mode, the pre-reactor will have a shorter runtime compared to pre- fully oxygen-saturated reactor (at start-up). Having front and rear ejectors allows for better oxygen absorption in the front ejector chemisorption reactor (pre-reactor). This option also provides the benefit of having an oxygen sensor between the front and rear expellers and having the option to change operation at exactly the point in time when the front is fully saturated or to keep it in the current for a while. more, if there is a disturbance of any nature in the process requiring a longer operation without exchange. Another benefit of this option is that the front ejector has to be significantly cooler than the main reactor; this is to prevent ethylene oxidation reaction from occurring. The rear ejector should preferably be warmer than the main ejector as most of the oxygen is removed and to remove trace oxygen a higher temperature is beneficial. Oxidation of the product stream in the expeller later than any significant point is not expected, as only traces of oxygen are present. Due to the operation as described above, when the pre-reactor (converter) changes to be a posterior expeller, it is still hot, which is very beneficial because the posterior, when changed posterior to be the front, is already cooled with ethylene product by the most efficient way, by direct contact of ethylene product with the catalyst surface.
    [085]The resulting product is then passed downstream for further separation if necessary. Separation requirements are minimized in the present reaction as the catalyst in the main reactor has a selectivity above 95% preferably above 98% and no or minimal by-products are produced in the oxygen removal step. The product can be sent directly to the polymerization plant or other facilities derived from ethylene, (such as ethylene glycol, acetic acid, vinyl acetate, etc.) as they may use ethylene of a lower purity, alternatively only CO , CO2 can be separated or just CO2 if necessary. However, as noted above, it is preferable to operate the prereactor in chemisorption or oxygen scavenger mode at one temperature to minimize additional generation of carbon dioxide, carbon monoxide, or both.
    [086] The present invention will now be illustrated by the following non-limiting examples. 1. Expulsion (Post-Removal) of Residual Oxygen from the Ethylene Product Gas Mix by the Periodic Redox Cycle
    [087]The expulsion of residual oxygen from the product mixture (emanated gas) was performed by cyclic periodic redox operation mode. In this case, a Mo1-V0.3-Nb0.2-Te0.1-OX catalyst or other Oxygen Storage Material (OSM) can be used in the two-step process. Step 1 provides for the reduction of the OSM layer by pure ethane at temperatures ~400°C, and Step 2 supports absorptive removal of residual O2 from the product mixture emanating from the pre-reduced layer which functions as an OSM at a reduced temperature. It was shown that the catalyst Mo1-V0.3-Nb0.2-Te0.1-OX itself served as a reasonably effective OSM at 300-400°Ca) Step 1: catalyst layer reduction Mo1-V0.3-Nb0.2 -Te0.1-OX by pure ethane
    [088] Measurements were made using a fresh sample [20% Mo-V-Te-Nb-Ox + 80% TiO2 (support)] prepared by mechanical methods (crushing and compaction/extrusion). In this test, the sample (2.0 cm3; 2.97 g, particle size 0.20.4 mm) was placed in a quartz reactor and heated to a specified temperature (375° and 400°C) in one flow of air, for 15 minutes, then the gas flow (900 cm3/h) was changed to pure ethane, and a sample of the emanated mixture was taken for analysis after a given time. After air reoxidation of the sample for 15 minutes, measurements were repeated several times with varying time intervals, and resulting product response curves were obtained (within 7.5 minutes). Figures 4, 5, and 6 demonstrate the time dependence of ethylene and CO2 formation rates as well as the selectivity of ethylene formation upon catalyst reduction by pure ethane at two different temperatures. The formation of CO curves are very similar to those observed for CO2 (Figure 5).
    [089] Figure 4 shows the profile (Dynamics) of ethylene formation reaction as a function of time at 375°C and 400°C after the gas flow changes [air ^ ethane] to a catalyst Mo1-V0.3- Nb0.2-Te0.1-OX with 80% TiO2 as support.
    [090] Figure 5 shows the reaction profile (Dynamics) of CO2 formation as a function of time at 375°C and 400°C after the gas flow changes [air ^ ethane]; for a Mo1-V0.3-Nb0.2-Te0.1-OX catalyst with 80% TiO2 as a support.
    [091] Figure 6 shows the selectivity of ethylene formation with respect to time at 375°C and 400°C after the gas flow changes [air ^ ethane]; by a Mo1-V0.3-Nb0.2-Te0.1-OX catalyst with 80% TiO2 as a support.
    [092]Thus, the catalyst reduction step by pure ethane at 380-400°C is accompanied by the formation of ethylene with a selectivity >92% (Figure 6).
    [093]The results obtained allow to calculate the total amount of “reactive” crosslinked oxygen in the catalyst that works as an OSM. The integration of the response curves (Figures 4 and 5) being produced at a constant ethane flow rate of 37.5 mmol/h allows one to assess the total amount of oxygen reacted during the catalyst reduction step. Taking into account that 1 g of the mixed oxide (Mo1-V0.3-Nb0.2-Te0.1-OX contains ~333 mg of oxygen, we can conclude that ~1.9% of this amount (~6.3 mg/g) ) can be removed from the active phase by reduction. Thus, the oxygen storage (absorption) capacity by the subsequent reoxidation step cannot exceed this number.(b)Step 2: Absorption of residual O2 from the mixture emanating from the layer pre-reduced which functions as an oxygen storage material at a reduced temperature
    [094] In this test, the catalyst sample after reduction by pure ethane at 400°C for 15 minutes was cooled to a given temperature (270°C) in the ethane stream, then a model gas product stream ([49 .5% by volume of C2H6 + 46.7% by volume of C2H4 + 3.8% by volume of O2 + traces of CO2]; 720 cm3/h) was changed, and the specimen of the emanating mixture was taken for analysis later of a given time. After subsequent sample reduction by ethane (15 minutes, 400°C), measurements were repeated several times with varying time intervals, and resulting product response curves were produced (within 5 minutes). The same test was repeated at 400°C for comparison. Figures 7 and 8 demonstrate the time dependence for total O2 removal, as well as the variation of CO and CO2 concentrations in the expulsion reaction flow product stream at 270°C and 400°C.
    [095] Figure 7 shows the profile (Dynamics) of the O2 removal reaction from the model gas mixture by the pre-reduced catalyst of (Mo1-V0.3-Nb0.2-Te0.1-OX supported in 80% of TiO2a 270°C and 400°C.
    [096] Figure 8 shows the profile (Dynamics) of CO2 reaction and CO formation after feeding the model gas mixture by the pre-reduced catalyst of (Mo1-V0.3-Nb0.2-Te0.1-OX supported in 80% TiO2a 270°C and 400°C
    [097] As can be seen, a very considerable formation of both CO and CO2 takes place at 400°C, ie not oxygen absorption, but instead the catalytic oxidation proceeds at a high temperature even in the pre-reduced catalyst . The situation changes after the temperature is reduced: at 270°C the absorptive removal of O2 becomes the main process, with a smaller contribution from the formation of both CO and CO2. INDUSTRIAL APPLICATION
    [098] The present invention seeks to improve the conversion and selectivity of oxidative dehydrogenation reactors by providing scavenger beds downstream of the reaction to remove residual oxygen from the product stream by interrupting the conversion of desired product(s) in products such as CO and CO2.
    权利要求:
    Claims (10)
    [0001]
    1. Process for catalytic oxidative dehydrogenation of one or more C2-4 alkanes comprising at least 2 pre-reactors (4 and 17) and one or more downstream main oxidative dehydrogenation reactors (8), CHARACTERIZED by the fact that it comprises: i) passing a feed stream comprising said one or more C2-4 alkanes through a first pre-reactor (4) containing a dehydrogenation catalyst that is saturated with reactive oxygen; ii) reacting the feed stream with the dehydrogenation catalyst; dehydrogenation that is saturated with reactive oxygen at a temperature of 300°C to 500°C and a pressure of 3.447 kPag to 689.47 kPag (0.5 to 100 psig) to produce a partially dehydrogenated stream comprising unreacted C2-4 alkanes ;iii) passing the partially dehydrogenated stream along with the additional oxygen feed to one or more downstream main oxidative dehydrogenation reactors (8); iv) oxidatively dehydrogenating the partially dehydrogenated stream to a temperature of 300°C to 500°C and a pressure of 3,447 kPag to 689.47 kPag (0.5 to 100 psig) to produce a product stream; v) removing the product stream from said one or more dehydrogenation reactors downstream major oxidatives (8) comprising C2-4 alkenes, unreacted C2-4 alkanes, corresponding unreacted oxygen and water vapor; vi) passing the product stream through a second prereactor (17) comprising a second catalyst of dehydrogenation that is depleted of reactive oxygen; vii) reacting the product stream with the dehydrogenation catalyst depleted of reactive oxygen at a temperature of 50°C to 270°C and a pressure of 3,447 kPag to 689.47 kPag (0, 5 to 100 psig) to complex oxygen by depleting oxygen from the product stream and regenerate the oxidative dehydrogenation catalyst by increasing saturation with reactive oxygen; and viii) recovering an oxygen depleted product stream; ix) continuing steps (i) to (viii) until: a) the pre-reactor (4) comprising the first dehydrogenation catalyst and through which the feed stream is being passed it is depleted of reactive oxygen or is more depleted of reactive oxygen than another pre-reactor (17); or b) the pre-reactor (17) comprising the second oxidative dehydrogenation catalyst through which the product stream is being passed is substantially saturated with reactive oxygen; x) when condition (a) is reached the feed stream is bypassed from first pre-reactor (4) to another pre-reactor (17) comprising a dehydrogenation catalyst that is saturated with reactive oxygen and the process continues from step (ii); xi) when condition (b) is reached the current of product is diverted from the second to the other pre-reactor comprising a second reactive oxygen depleted dehydrogenation catalyst and the process continues from step (viii).
    [0002]
    2. Process according to claim 1, CHARACTERIZED by the fact that the oxidative dehydrogenation catalyst in any reactor is independently selected from: 1) 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; e11) 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 0.3 , d is a number to satisfy the valence state of the mixed oxide catalyst.
    [0003]
    3. Process according to claim 2, CHARACTERIZED by the fact that said one or more downstream main oxidative dehydrogenation reactors (8) are operated at a gas hourly space velocity (GHSV) of 500 to 30,000 h-1 .
    [0004]
    4. Process according to claim 3, CHARACTERIZED by the fact that the catalyst in the main reactor is supported on an inert metal oxide support having a surface area of less than 5 m2/g.
    [0005]
    5. Process according to claim 4, CHARACTERIZED by the fact that the pre-reactors are fixed bed reactors and the oxidative dehydrogenation catalyst is supported on an inert metal oxide support having a surface area of not less than 100 m2/g.
    [0006]
    6. Process according to claim 5, CHARACTERIZED by the fact that said one or more downstream main oxidative dehydrogenation reactors (8) are selected from fixed bed reactors (30, 41), fluidized bed reactors or boiling and ceramic membrane reactors.
    [0007]
    7. Process according to claim 6, CHARACTERIZED by the fact that said one or more C2-4 alkanes is ethane.
    [0008]
    8. Process according to claim 7, CHARACTERIZED by the fact that the catalyst in said pre-reactors is of formula (ii) and in which: a is from 0.90 to 1.10; b is from 0.25 to 0.3;c is from 0.1 to 0.3;e is from 0.1 to 0.2; d is a number to satisfy the valence state of the mixed oxide catalyst.
    [0009]
    9. Process according to claim 1, CHARACTERIZED by the fact that in step (vi), the product stream from said downstream main oxidative dehydrogenation reactors (8) is sequentially passed through two or more pre-reactors in the which the first of these sequential pre-reactors has a higher amount of reactive oxygen in the oxidative dehydrogenation catalyst than in the subsequent pre-reactors.
    [0010]
    10. Process according to claim 1, CHARACTERIZED by the fact that it further comprises a third pre-reactor comprising an oxidative dehydrogenation catalyst, wherein the third pre-reactor can:i) accept the feed stream when the catalyst oxidative dehydrogenation in the third pre-reactor is enriched with oxygen; or ii) accept the product stream when the oxidative dehydrogenation catalyst in the third pre-reactor is depleted of oxygen.
    类似技术:
    公开号 | 公开日 | 专利标题
    BR112016010786B1|2021-08-17|PROCESS FOR CATALYTIC OXIDATIVE DEHYDROGENATION OF ONE OR MORE C2-4 ALKANS
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    US10343957B2|2019-07-09|Complex comprising ODH unit with integrated oxygen separation module
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    同族专利:
    公开号 | 公开日
    CA2833822C|2020-08-04|
    KR102304410B1|2021-09-23|
    US20170081261A1|2017-03-23|
    NO3071539T3|2018-03-31|
    CA2833822A1|2015-05-21|
    ES2655914T3|2018-02-22|
    MX2016006186A|2016-08-08|
    JP2019194245A|2019-11-07|
    US9550709B2|2017-01-24|
    EP3071539A1|2016-09-28|
    JP2016539124A|2016-12-15|
    WO2015075597A1|2015-05-28|
    BR112016010786A2|2017-08-08|
    KR20160090322A|2016-07-29|
    EP3071539B1|2017-11-01|
    CN105722810A|2016-06-29|
    US20150141727A1|2015-05-21|
    CN105722810B|2018-02-02|
    JP6557230B2|2019-08-07|
    US9630893B2|2017-04-25|
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    法律状态:
    2020-01-14| B06U| Preliminary requirement: requests with searches performed by other patent offices: procedure suspended [chapter 6.21 patent gazette]|
    2020-09-29| B06A| Patent application procedure suspended [chapter 6.1 patent gazette]|
    2021-07-27| B09A| Decision: intention to grant [chapter 9.1 patent gazette]|
    2021-08-17| B16A| Patent or certificate of addition of invention granted [chapter 16.1 patent gazette]|Free format text: PRAZO DE VALIDADE: 20 (VINTE) ANOS CONTADOS A PARTIR DE 10/11/2014, OBSERVADAS AS CONDICOES LEGAIS. |
    优先权:
    申请号 | 申请日 | 专利标题
    CA2.833,822|2013-11-21|
    CA2833822A|CA2833822C|2013-11-21|2013-11-21|Inherently safe odh operation|
    PCT/IB2014/065930|WO2015075597A1|2013-11-21|2014-11-10|Inherently safe odh operation|
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