Methods for converting alkanes to alkenes

Abstract

A method for converting alkanes to alkenes includes contacting a feed stream comprising alkanes with a dehydrogenation catalyst in a reaction zone, the dehydrogenation catalyst comprising molybdenum, vanadium, at least 0.1 wt.% terbium, and at least 0.5 wt.% antimony, wherein the weight percent is based on a total weight of the dehydrogenation catalyst, and converting at least a portion of the alkanes to alkenes, thereby yielding a product stream comprising alkanes, alkenes, and steam.

Description

METHODS FOR CONVERTING ALKANES TO ALKENESCROSS-REFERENCE TO RELATED APPLICATION

[0001] This application claims priority to U.S. Provisional Application Serial No. 63 / 665,804 filed June 28, 2024, the entire disclosure of which is hereby incorporated by reference.TECHNICAL FIELD

[0002] The present disclosure relates to methods for converting various alkanes to alkenes.BACKGROUND

[0003] Ethene (commonly known as ethylene or C2H4) is a commodity chemical widely used to produce plastics such as polyethylene, with an annual production exceeding 150 million metric tons. Steam cracking of ethane or naphtha is the dominant pathway for C2H4 production. This endothermic process operated between 700 and 950 °C is highly energy-intensive, which results in global annual CO2 emissions greater than 260 million tonnes. Unfortunately, this technology is already highly optimized for energy efficiency due to its maturity. Thus, it is very challenging to make incremental gains to further limit emissions in this process in the face of growing demand for C2H4. Thus, innovative solutions to produce C2H4 that enable the chemical industry to scale down its CO2 emissions would be desirable.SUMMARY

[0004] Embodiments of the present disclosure address these and other needs by providing methods for converting alkanes to alkenes including contacting a feed stream comprising alkanes with a dehydrogenation catalyst in a reaction zone, the dehydrogenation catalyst comprising molybdenum, vanadium, at least 0.1 wt.% terbium, and at least 0.5 wt.% antimony, wherein the weight percent is based on a total weight of the dehydrogenation catalyst , and converting at least a portion of the alkanes to alkenes, thereby yielding a product stream comprising alkanes, alkenes, and steam.

[0005] Additional features and advantages will be set forth in the detailed description that follows and, in part, will be readily apparent to those skilled in the art from that description orrecognized by practicing the embodiments described herein, including the detailed description which follows in addition to the claims.

[0006] It is to be understood that both the foregoing general description and the following detailed description describe various embodiments and are intended to provide an overview or framework for understanding the nature and character of the claimed subject matter.BRIEF DESCRIPTION OF THE DRAWINGS

[0007] FIG. 1 graphically depicts the production rate of C2H4 with respect to the depletion of oxygen.

[0008] FIG. 2 graphically depicts the ratio of cumulative C2H4 produced to catalyst mass ratio with respect to the depletion of oxygen.

[0009] FIG. 3 graphically depicts the conversion, selectivity, and yield plots of the catalyst of the examples according to embodiments disclosed herein with respect to time.

[0010] FIG. 4 graphically depicts C2H4 yield versus the nominal atomic ratio of terbium and antimony according to examples according to embodiments disclosed herein.

[0011] FIG. 5 depicts the X-ray diffraction (XRD) of various examples according to embodiments disclosed herein.DETAILED DESCRIPTION

[0012] Reference will now be made in detail to embodiments of methods for converting alkanes to alkenes.

[0013] As used herein, “dehydrogenation” refers to a chemical process by which hydrogen is chemically removed from a chemical compound. For example, ethane may undergo dehydrogenation to be converted to ethylene. As used herein, “dehydrogenation catalyst(s)” refers to any substance that increases the rate of a dehydrogenation reaction without itself undergoing any permanent chemical change upon either simultaneous exposure with alkane and oxygen (O2) or reacting with alkane and O2 in temporally or spatially separated step(s).

[0014] As used herein, “alkane(s)” refers to any series of hydrocarbon molecules that consist of carbon-carbon single bonds and where the carbon structure is saturated with hydrogen. Ethane, propane, and butane are examples of alkanes. As used herein, “alkene(s)” refers to any series of hydrocarbon molecules, where at least two of the carbon atoms are not saturated with hydrogen and share a double bond. Ethylene, propylene, 1 -butene, / ra / / .s-2-butene, and c / .s-2-butene are examples of alkenes. Dienes are hydrocarbons where at least two sets of two carbon molecules, which may or may not be adjacent to each other, are not saturated with hydrogen and share a double bond. Thus, dienes may include a combination of alkenes. Accordingly, unless explicitly stated otherwise, “alkenes” as used herein includes dienes.

[0015] An alternative means to using steam cracking to produce C2H4 is the catalytic oxidative dehydrogenation (ODH) process which applies a heterogeneous catalyst for the exothermic reaction between C2H6 and O2 operated at a far lower temperature (350 - 500 °C) to selectively form C2H4. The lower temperature limits the formation of compounds with more than three carbon atoms, thereby reducing risks of fouling and additional expenses related to their separation from C2H4. Co-feeding O2 also minimizes carbon deposition, which is unavoidable and costly for steam cracker operation. A disadvantage of ODH is that it typically co-produces acetic acid, which has a relatively small market size. Another disadvantage is the safety risks of combining C2H6 and O2 for an exothermic reaction in a reactor. Apart from that, the process requires high purity O2, whose production requires additional energy input.

[0016] The method of the present disclosure solves these and other problems by operating the oxidative dehydrogenation reaction in a transient manner by feeding C2H6 to a lattice oxygen-rich dehydrogenation catalyst in the absence of a gaseous O2 co-feed to produce H2O and C2H4, resulting in the loss of lattice oxygen from the dehydrogenation catalyst. In a spatially or temporally separated step, the lost oxygen from the dehydrogenation catalyst is regenerated by reoxidizing the dehydrogenation catalyst in air. This process is termed lattice oxidative dehydrogenation (LODh) and operated at a temperature similar to ODH. By avoiding co-feeding O2 and C2H6, safety risks are circumvented. By using air to regenerate the dehydrogenation catalyst, the need for separate O2 production is avoided.

[0017] Molybdenum vanadium oxide catalysts, when used for LODh of ethane, typically show very high initial ethane conversion (>30%), which generally drops to a very low value (<10%)with time in the absence of co-fed O2. This is typically caused by depletion of lattice oxygen and, possibly, increasingly slower diffusion oxygen atoms from the catalyst bulk to the surface.

[0018] Adding terbium (Tb) to molybdenum vanadium oxide catalysts significantly impacts the oxygen release rate of the resulting dehydrogenation catalyst. The addition appears to sustain C2H4 production for a longer time. Without being bound by theory, it is believed that C2H4 production is sustained because the diffusion of lattice oxygen to the dehydrogenation catalyst surface is less impacted by the degree of oxygen depletion (reduction) in the dehydrogenation catalyst. The ability of the material to sustain oxygen transport to the surface has the potential to improve reactor productivity, maintain product quality, and simplify reactor design and downstream purification.

[0019] In one or more embodiments, a method for converting alkanes to alkenes may comprise contacting a feed stream comprising alkanes with a dehydrogenation catalyst in a reaction zone. In embodiments, the feed stream may comprise C2-C4 alkanes. In embodiments, the feed stream may comprise ethane, propane, butanes, or combinations thereof. In embodiments, the feed stream may be contacted with the dehydrogenation catalyst for a controlled time of exposure. In embodiments, the controlled time of exposure may be selected based on the desired dehydrogenation catalyst to feed stream mass to mass ratio.

[0020] In one or more embodiments, the dehydrogenation catalyst includes molybdenum (Mo), vanadium (V), and terbium (Tb). It has been unexpectedly found that dehydrogenation catalysts that include terbium, in addition to molybdenum and vanadium, significantly impacts the oxygen release rate of the dehydrogenation catalyst. This addition of terbium appears to sustain alkene production for a longer time suggesting that the diffusion of lattice oxygen to the dehydrogenation catalyst surface is less impacted by the degree of oxygen depletion (reduction) in the dehydrogenation catalyst. The ability of the dehydrogenation catalyst to sustain oxygen transport to the dehydrogenation catalyst surface has the potential to improve reactor productivity, maintain product quality, and simplify reactor design and downstream purification.

[0021] In embodiments, the disclosed dehydrogenation catalyst may be substantially free of tellurium (Te), meaning that the dehydrogenation catalyst comprises less than 0.001 wt.% Te or a Te / Mo ratio below 0.01.

[0022] The dehydrogenation catalyst may include 33 wt.% to 66 wt.% molybdenum, such as from 33 wt.% to 66 wt.%, from 35 wt.% to 66 wt.%, from 40 wt.% to 66 wt.%, from 45 wt.% to 66 wt.%, from 50 wt.% to 66 wt.%, from 33 wt.% to 65 wt.%, from 35 wt.% to 65 wt.%, from 40 wt.% to 65 wt.%, from 45 wt.% to 65 wt.%, from 50 wt.% to 65 wt.%, from 33 wt.% to 60 wt.%, from 35 wt.% to 60 wt.%, from 40 wt.% to 60 wt.%, from 45 wt.% to 60 wt.%, from 50 wt.% to 60 wt.%, from 33 wt.% to 55 wt.%, from 35 wt.% to 55 wt.%, from 40 wt.% to 55 wt.%, from 45 wt.% to 55 wt.%, from 50 wt.% to 55 wt.%, from 33 wt.% to 50 wt.%, from 35 wt.% to 50 wt.%, from 40 wt.% to 50 wt.%, or from 45 wt.% to 50 wt.% molybdenum, wherein the weight percent is based on a total weight of the dehydrogenation catalyst.

[0023] The dehydrogenation catalyst may include 2 wt.% to 16 wt.% vanadium, such as from 2 wt.% to 16 wt.%, from 4 wt.% to 16 wt.%, from 5 wt.% to 16 wt.%, from 6 wt.% to 16 wt.%, from 8 wt.% to 16 wt.%, from 9 wt.% to 16 wt.%, from 10 wt.% to 16 wt.%, from 2 wt.% to 15 wt.%, from 4 wt.% to 15 wt.%, from 5 wt.% to 15 wt.%, from 6 wt.% to 15 wt.%, from 8 wt.% to 15 wt.%, from 9 wt.% to 15 wt.%, from 10 wt.% to 15 wt.%, from 2 wt.% to 12 wt.%, from 4 wt.% to 12 wt.%, from 5 wt.% to 12 wt.%, from 6 wt.% to 12 wt.%, from 8 wt.% to 12 wt.%, from 9 wt.% to 12 wt.%, from 10 wt.% to 12 wt.%, from 2 wt.% to 10 wt.%, from 4 wt.% to 10 wt.%, from 5 wt.% to 10 wt.%, from 6 wt.% to 10 wt.%, from 8 wt.% to 10 wt.%, from 9 wt.% to 10 wt.%, from 2 wt.% to 9 wt.%, from 4 wt.% to 9 wt.%, from 5 wt.% to 9 wt.%, from 6 wt.% to 9 wt.%, or from 8 wt.% to 9 wt.% vanadium, wherein the weight percent is based on a total weight of the dehydrogenation catalyst.

[0024] The dehydrogenation catalyst may include 0.1 wt.% to 19 wt.% terbium, such as from 0.1 wt.% to 19 wt.%, from 0.5 wt.% to 19 wt.%, from 1 wt.% to 19 wt.%, from 2 wt.% to 19 wt.%, from 4 wt.% to 19 wt.%, from 5 wt.% to 19 wt.%, from 6 wt.% to 19 wt.%, from 8 wt.% to 19 wt.%, from 10 wt.% to 19 wt.%, from 0.1 wt.% to 18 wt.%, from 0.5 wt.% to 18 wt.%, from 1 wt.% to 18 wt.%, from 2 wt.% to 18 wt.%, from 4 wt.% to 18 wt.%, from 5 wt.% to 18 wt.%, from 6 wt.% to 18 wt.%, from 8 wt.% to 18 wt.%, from 10 wt.% to 18 wt.%, terbium, from 0.1 wt.% to 16 wt.%, from 0.5 wt.% to 16 wt.%, from 1 wt.% to 16 wt.%, from 2 wt.% to 16 wt.%, from 4 wt.% to 16 wt.%, from 5 wt.% to 16 wt.%, from 6 wt.% to 16 wt.%, from 8 wt.% to 16 wt.%, from 10 wt.% to 16 wt.%, from 0.1 wt.% to 15 wt.%, from 0.5 wt.% to 15 wt.%, from 1 wt.% to 15 wt.%, from 2 wt.% to 15 wt.%, from 4 wt.% to 15 wt.%, from 5 wt.% to 15 wt.%, from 6 wt.% to 15 wt.%, from 8 wt.% to 15 wt.%, from 10 wt.% to 15 wt.%, from 0.1 wt.% to 14 wt.%, from 0.5wt.% to 14 wt.%, from 1 wt.% to 14 wt.%, from 2 wt.% to 14 wt.%, from 4 wt.% to 14 wt.%, from5 wt.% to 14 wt.%, from 6 wt.% to 14 wt.%, from 8 wt.% to 14 wt.%, from 10 wt.% to 14 wt.%, from 0.1 wt.% to 12 wt.%, from 0.5 wt.% to 12 wt.%, from 1 wt.% to 12 wt.%, from 2 wt.% to 12 wt.%, from 4 wt.% to 12 wt.%, from 5 wt.% to 12 wt.%, from 6 wt.% to 12 wt.%, from 8 wt.% to 12 wt.%, from 10 wt.% to 12 wt.%, from 0.1 wt.% to 19 wt.%, from 0.5 wt.% to 19 wt.%, from 1 wt.% to 19 wt.%, from 2 wt.% to 19 wt.%, from 4 wt.% to 19 wt.%, from 5 wt.% to 19 wt.%, from6 wt.% to 19 wt.%, or from 8 wt.% to 19 wt.% terbium wherein the weight percent is based on a total weight of the dehydrogenation catalyst.

[0025] In some embodiments, the dehydrogenation catalyst may further include antimony (Sb). It has been unexpectedly found that dehydrogenation catalyst that include molybdenum and vanadium as well as terbium and antimony demonstrate a capability to sustain C2H4 production rates. Without being bound by theory, it is believed that the capability to sustain C2H4 production rates caused by the combination of terbium and antimony is caused by faster diffusion of lattice oxygen to the surface of a dehydrogenation catalyst.

[0026] The dehydrogenation catalyst may include 0.5 wt.% to 14 wt.% antimony, such as from0.5 wt.% to 14 wt.%, from 1 wt.% to 14 wt.%, from 2 wt.% to 14 wt.%, from 4 wt.% to 14 wt.%, from 5 wt.% to 14 wt.%, from 6 wt.% to 14 wt.%, from 8 wt.% to 14 wt.%, from 0.5 wt.% to 12 wt.%, from 1 wt.% to 12 wt.%, from 2 wt.% to 12 wt.%, from 4 wt.% to 12 wt.%, from 5 wt.% to12 wt.%, from 6 wt.% to 12 wt.%, from 8 wt.% to 12 wt.%, from 0.5 wt.% to 10 wt.%, from 1 wt.% to 10 wt.%, from 2 wt.% to 10 wt.%, from 4 wt.% to 10 wt.%, from 5 wt.% to 10 wt.%, from6 wt.% to 10 wt.%, from 8 wt.% to 10 wt.%, from 0.5 wt.% to 8 wt.%, from 1 wt.% to 8 wt.%, from 2 wt.% to 8 wt.%, from 4 wt.% to 8 wt.%, from 5 wt.% to 8 wt.%, or from 6 wt.% to 8 wt.% antimony, wherein the weight percent is based on a total weight of the dehydrogenation catalyst.

[0027] The dehydrogenation catalyst may include an atomic ratio of vanadium to molybdenum of from 0.1 : 1 to 0.6: 1, such as from 0.2: 1 to 0.6: 1, from 0.3: 1 to 0.6: 1, from 0.4: 1 to 0.6: 1, from 0.5: 1 to 0.6: 1, from 0.1 : 1 to 0.5: 1, from 0.2: 1 to 0.5:1, from 0.3: 1 to 0.5: 1, from 0.4: 1 to 0.5: 1, from 0.1 : 1 to 0.4: 1, from 0.2: 1 to 0.4: 1, from 0.3: 1 to 0.4: 1, from 0.1 : 1 to 0.3: 1, from 0.2: 1 to 0.3: 1, or from 0.1 : 1 to 0.2: 1.

[0028] The dehydrogenation catalyst may include an atomic ratio of terbium to molybdenum of from 0.002:1 to 0.2:1, from 0.005:1 to 0.2:1, from 0.01:1 to 0.2:1, from 0.05:1 to 0.2:1, from 0.1:1 to 0.2:1, from 0.002:1 to 0.1:1, from 0.005:1 to 0.1:1, from 0.01:1 to 0.1:1, from 0.05:1 to 0.1:1, from 0.002:1 to 0.05:1, from 0.005:1 to 0.05:1, from 0.01:1 to 0.05:1, from 0.002:1 to 0.01:1, from 0.005:1 to 0.01:1, or from 0.002:1 to 0.005:1.

[0029] The dehydrogenation catalyst may include an atomic ratio of antimony to molybdenum of from 0:1 to 0.2:1, from 0.001:1 to 0.2:1, from 0.005:1 to 0.2:1, from 0.01: 1 to 0.2:1, from 0.05:1 to 0.2:1, or from 0.1:1 to 0.2:1.

[0030] In some embodiments, molybdenum, vanadium, and terbium may be present in the dehydrogenation catalyst as oxides, i.e. as molybdenum oxide, vanadium oxide, and terbium oxide. In some embodiments, the antimony may be present in the dehydrogenation catalyst as an oxide, i.e. as antimony oxide. In some embodiments, niobium and tantalum may be present in the dehydrogenation catalyst as oxides, i.e. as niobium oxide and tantalum oxide.

[0031] In various embodiments, the dehydrogenation catalyst may have an atomic ratio of terbium to antimony from 0.05:1 to 20:1. The dehydrogenation catalyst may have an atomic ratio of terbium to antimony from 0.05:1 to 20:1, from 0.1:1 to 20:1, from 0.5:1 to 20:1, from 1:1 to 20:1, from 2:1 to 20:1, from 5:1 to 20:1, from 10:1 to 20:1, from 15:1 to 20:1, from 0.05:1 to 15:1, from 0.1:1 to 15:1, from 0.5:1 to 15:1, from 1:1 to 15:1, from 2:1 to 15:1, from 5:1 to 15:1, from 10:1 to 15:1, from 0.05:1 to 10:1, from 0.1:1 to 10:1, from 0.5:1 to 10:1, from 1:1 to 10:1, from 2:1 to 10:1, from 5:1 to 10:1, from 0.05:1 to 5:1, from 0.1:1 to 5:1, from 0.5:1 to 5:1, from 1:1 to 5:1, from 2:1 to 5:1, from 0.05:1 to 2:1, from 0.1:1 to 2:1, from 0.5:1 to 2:1, from 1:1 to 2:1, from 0.05:1 to 1:1, from 0.1:1 to 1:1, from 0.5:1 to 1:1, from 0.05:1 to 0.5:1, from 0.1:1 to 0.5:1, or from 0.05:1 to 0.1:1.

[0032] In some embodiments, the dehydrogenation catalyst may have an atomic ratio of antimony to molybdenum from 0.01:1 to 0.2: 1. The dehydrogenation catalyst may have an atomic ratio of from 0.01:1 to 0.2:1, from 0.05:1 to 0.2:1, from 0.1:1 to 0.2:1, from 0.01:1 to 0.1:1, from 0.05:1 to 0.1:1, or from 0.01:1 to 0.05:1.

[0033] In one or more embodiments, the dehydrogenation catalyst may further include niobium (Nb), tantalum (Ta), or combinations thereof. The dehydrogenation catalyst may include at least 0.04 wt.% of niobium, tantalum, or combinations thereof. The dehydrogenation catalyst may include from 0.04 wt.% to 11 wt.%, 0.04 wt.% to 10 wt.%, 0.04 wt.% to 8 wt.%, 0.04 wt.% to 6 wt.%, 0.04 wt.% to 5 wt.%, from 0.04 wt.% to 2 wt.%, from 0.04 wt.% to 1 wt.%, from 0.04 wt.% to 0.5 wt.%, from 0.04 wt.% to 0.1 wt.%, or from 0.04 wt.% to 0.05 wt.%, niobium, tantalum, or combinations thereof.

[0034] In various embodiments, the dehydrogenation may include an atomic ratio of niobium to molybdenum of from 0: 1 to 0.1 : 1, from 0: 1 to 0.05: 1, from 0: 1 to 0.01 : 1, from 0.01 : 1 to 0.1 : 1, from 0.01 :1 to 0.05: 1, or from 0.05: 1 to 0.1: 1. The dehydrogenation catalyst may have an atomic ratio of tantalum to molybdenum of from 0: 1 to 0.1 : 1, from 0: 1 to 0.05: 1, from 0: 1 to 0.01 : 1, from 0.01 : 1 to 0.1 : 1, from 0.01 : 1 to 0.05: 1, or from 0.05: 1 to 0.1 : 1. The dehydrogenation catalyst may have a sum of the atomic ratios of niobium to molybdenum and tantalum to molybdenum of from 0.001 : 1 to 0.1 : 1, from 0.005: 1 to 0.1 :1, from 0.01 :1 to 0.1 :1, from 0.05:1 to 0.1 :1, from 0.001 :1 to 0.05: 1, from 0.005: 1 to 0.05: 1, from 0.01 : 1 to 0.05: 1, from 0.001 : 1 to 0.01 : 1, from 0.005: 1 to 0.01 : 1, or from 0.001 : 1 to 0.005: 1.

[0035] In some embodiments, the dehydrogenation catalyst may include a crystal structure with a Pba2-32 space group. Without being bound by theory, it is believed that a crystal structure with a Pba2-32 space group is believed to be an active phase of the dehydrogenation catalyst.

[0036] In one or more embodiments, the reaction zone may be a zone inside a reactor adapted to allow the feed stream to be contacted with the dehydrogenation catalyst. In one or more embodiments, the reactor may be a fixed-bed reactor, including but not limited to a dual tube fixed-bed reactor. In embodiments, the reactor may be a circulating fluidized bed reactor. In embodiments, the reactor may be two or more reactors in series or parallel, and each reactor in series or parallel may be the same type of reactor as other reactors in the series, or may be a different type of reactor from other reactors in the series.

[0037] In one or more embodiments, a method for converting alkanes to alkenes may comprise converting at least a portion of the alkanes to alkenes, thereby yielding a product stream comprising alkanes, alkenes, and steam. In embodiments, the product stream may compriseethylene, propylene, butylene, hydrogen, steam, or combinations thereof. In embodiments, the product stream may comprise ethane, propane, butane, ethylene, propylene, butylene, hydrogen, steam, or combinations thereof. In one or more embodiments, at least a portion of the hydrogen in the product stream is combusted and yields water. In embodiments, the water may be in the form of steam. In embodiments, the steam may comprise gaseous water, liquid water, aerosolized water, or combinations thereof. Because of this hydrogen combustion, water — such as steam — may be present in the reaction zone during dehydrogenation of the alkanes in the feed stream.

[0038] As mentioned above, many conventional alkane dehydrogenation processes, such as oxidative dehydrogenation processes, require the use of gaseous oxidants, such as oxygen, air, carbon dioxide, or nitrogen oxides, in the feed stream or as a co-feed. The term “gaseous oxidant(s)” may refer to a substance or substances other than water that may oxidize hydrogen. However, in one or more embodiments of the present disclosure, the dehydrogenation catalyst maintains the conversion of alkanes to alkenes without the presence of a gaseous oxidant in the feed stream or as a co-feed. In some embodiments, the dehydrogenation catalyst maintains the conversion of alkanes to alkenes with the presence of only a small amount of a gaseous oxidant in the feed stream or as a co-feed. In some embodiments, the dehydrogenation catalyst maintains the conversion of alkanes to alkenes with the presence of less than 5 v.%, less than 4 v.%, less than 3 v.%, less than 2 v.%, less than 1 v.%, less than 0.5 v.%, less than 0.25 v.%, or less than 0.1 v.% gaseous oxidant in the feed stream or as a co-feed.

[0039] In further embodiments, the feed stream may consist essentially of alkanes, which herein means the feed stream may comprise at least 95 vol.% alkanes, such as at least 97 vol.% alkanes or at least 99 vol.% alkanes. In one or more embodiments, the feed stream may comprise greater than or equal to 40 vol.% alkanes, such as greater than or equal to 50 vol.% alkanes, as greater than or equal to 70 vol.% alkanes, or greater than or equal to 90 vol.% alkanes. In one or more embodiments, the feed stream may comprise greater than or equal to 40 vol.% alkanes and less than or equal to 60 vol.% oxygen, steam and / or inert gas, such as greater than or equal to 50 vol.% alkanes and less than or equal to 50 vol.% oxygen, steam and / or inert gas, greater than or equal to 70 vol.% alkanes and less than or equal to 30 vol.% oxygen, steam and / or inert gas, or greater than or equal to 90 vol.% alkanes and less than or equal to 10 vol.% oxygen, steam and / or inert gas.

[0040] In one or more embodiments, the feed stream may be substantially free of a gaseous oxidant, meaning that the feed stream comprises less than 1 vol.% oxidants, such as less than 0.5 vol.% oxidants, less than 0.1 vol.% oxidants, less than 0.05 vol.% oxidants, or less than 0.01 vol.% oxidants. In embodiments, the gaseous oxidant may be O2, NO, NO2, H2O, CO2, N2O, O3, or a combination thereof.

[0041] In one or more embodiments, the dehydrogenation catalyst and the alkanes in the feed stream have a mass to mass ratio that is from 5:1 to 500:1. In embodiments, the dehydrogenation catalyst and the feed stream have a mass to mass ratio that is from 5: 1 to 500:1, from 10: 1 to 500: 1, from 25:1 to 500:1, from 50:1 to 500:1, from 75:1 to 500:1, from 100:1 to 500:1, from 150:1 to 500:1, from 200:1 to 500:1, from 300:1 to 500:1, from 400:1 to 500:1, from 5:1 to 400:1, from 10:1 to 400:1, from 25:1 to 400:1, from 50:1 to 400:1, from 75:1 to 400:1, from 100:1 to 400:1, from 150:1 to 400:1, from 200:1 to 400:1, from 300:1 to 400:1, from 5:1 to 300:1, from 10:1 to 300:1, from 25:1 to 300:1, from 50:1 to 300:1, from 75:1 to 300:1, from 100:1 to 300:1, from 150:1 to 300:1, from 200:1 to 300:1, from 5:1 to 200:1, from 10:1 to 200:1, from 25:1 to 200:1, from 50:1 to 200:1, from 75:1 to 200:1, from 100:1 to 200:1, from 150:1 to 200:1, from 5:1 to 150:1, from 10:1 to 150:1, from 25:1 to 150:1, from 50:1 to 150:1, from 75:1 to 150:1, from 100:1 to 150:1, from 5:1 to 100:1, from 10:1 to 100:1, from 25:1 to 100:1, from 50:1 to 100:1, from 75:1 to 100:1, from 5:1 to 75:1, from 10:1 to 75:1, from 25:1 to 75:1, from 50:1 to 75:1, from 5:1 to 50:1, from 10:1 to 50:1, from 25:1 to 50:1, from 5:1 to 25:1, from 10:1 to 25:1, or from 5:1 to 10: 1. In one or more embodiments, the feed stream may enter the reactor at a temperature of from 300 °C to 600 °C, a pressure of from 0 bar(g) to 20 bar(g), and a weight hourly space velocity (WHSV) of from 0.1 / hr to 10 / hr. In further embodiments, the feed stream may enter the reactor at, for example, a temperature of from 300 °C to 500 °C, from 400 °C to 550 °C, from 450 °C to 600 °C, or from 500 °C to 600 °C; a pressure of from 0 bar(g) to 5 bar(g), from 0 bar(g) to 10 bar(g), from 0 bar(g) to 15 bar(g), from 5 bar(g) to 15 bar(g), from 10 bar(g) to 15 bar(g), from 10 bar (g) to 20 bar(g), or from 15 bar(g) to 20 bar(g); and a weight hourly space velocity (WHSV) of from 0.1 / hr to 3 / hr, from 0.1 / hr to 5 / hr, from 0.1 / hr to 7 / hr, from 3 / hr to 5 / hr, from 3 / hr to 7 / hr, from 5 / hr to 7 / hr, from 5 / hr to 10 / hr, or from 7 / hr to 10 / hr.

[0042] In one or more embodiments, the reactor may be substantially free of a gaseous oxidant.

[0043] In one or more embodiments, an oxygen stream may be added to a reaction zone. The concentration of oxygen in the oxygen stream is not particularly limited. For example, the oxygen concentration in the oxygen stream may be from 0.1 vol% to 99.9 vol%, such as from 5.0 vol% to 95.0 vol%, from 10.0 vol% to 90.0 vol%, from 15.0 vol% to 85.0 vol%, from 20.0 vol% to 80.0 vol%, from 25.0 vol% to 75.0 vol%, from 30.0 vol% to 70.0 vol%, from 35.0 vol% to 65.0 vol%, from 40.0 vol% to 60.0 vol%, or from 45.0 vol% to 55.0 vol%. In one or more embodiments, the concentration of oxygen in the oxygen stream is relatively low, such as from 0.1 vol% to 5.0 vol%, from 0.2 vol% to 5.0 vol%, from 0.5 vol% to 5.0 vol%, from 0.8 vol% to 5.0 vol%, from 1.0 vol% to 5.0 vol%, from 1.2 vol% to 5.0 vol%, from 1.5 vol% to 5.0 vol%, from 1.8 vol% to 5.0 vol%, from 2.0 vol% to 5.0 vol%, from 2.2 vol% to 5.0 vol%, from 2.5 vol% to 5.0 vol%, from 2.8 vol% to 5.0 vol%, from 3.0 vol% to 5.0 vol%, from 3.2 vol% to 5.0 vol%, from 3.5 vol% to 5.0 vol%, from 3.8 vol% to 5.0 vol%, from 4.0 vol% to 5.0 vol%, from 4.2 vol% to 5.0 vol%, from 4.5 vol% to 5.0 vol%, or from 4.8 vol% to 5.0 vol%. In embodiments, the oxygen stream may be air, which generally has an oxygen concentration of about 21.0 vol%.

[0044] The oxygen stream may, in embodiments, be added to the reaction zone sequentially to the feed stream, such that the feed stream and the oxygen stream are not added to the reaction zone at the same time. It should be understood that in embodiments, the oxygen stream may be added at various points within the reaction process. This may be accomplished by introducing the oxygen stream into the reaction zone at different locations within the reaction zone and / or introducing the oxygen stream at different time periods while the reaction is taking place.

[0045] In one or more embodiments, the oxygen stream is added to the reaction zone simultaneously to the feed stream. In such embodiments, the volume ratio of oxygen (in the oxygen stream) to alkanes (in the feed stream) in the reaction zone is from greater 0.0 to 1.0, from 0.1 to 1.0, from 0.2 to 1.0, from 0.3 to 1.0, from 0.4 to 1.0, from 0.5 to 1.0, from 0.6 to 1.0, from 0.7 to 1.0, from 0.8 to 1.0, from 0.9 to 1.0, or from greater than 0.0 to 0.2, from 0.0 to 0.3, from 0.0 to 0.4, from 0.0 to 0.5, from 0.0 to 0.6, from 0.0 to 0.7, from 0.0 to 0.8, or from 0.0 to 0.9.

[0046] The reactor is not particularly limited and any type of reactor allowing for cyclic or continuous operation of the dehydrogenation process may be used in embodiments. In embodiments, the reactor may be a fixed bed reactor, a fluidized bed reactor, a moving bed reactor, a bubbling bed reactor, circulating fluidized reactor, reverse flow reactor, or an ebullated(bubbling) bed reactor. In embodiments, a reaction zone can be all or apportion of the reactor. The reaction zone is not particularly limited to a single reactor and can consist of multiple reactors in either series or parallel configuration.

[0047] Reference will now be made in detail to embodiments of dehydrogenation catalysts that are capable of catalyzing the alkane-alkene conversion substantially in the absence of a gaseous oxidant.

[0048] Using dehydrogenation catalysts disclosed herein convert alkanes to alkenes, and particularly, ethane to ethylene, substantially in the absence of a gaseous oxidant at a temperature similar to or lower than conventional oxidative dehydrogenation or “ODH” processes. These oxidative dehydrogenation catalysts comprise lattice oxygen species. Without being bound by any particular theory, lattice oxygen can activate alkanes, thus this particular dehydrogenation process is termed lattice oxidative dehydrogenation or “LODh”.

[0049] It has been observed that previously known molybdenum-vanadium oxide based catalysts may be capable of catalyzing both ODH and LODh processes. However, when known molybdenum vanadium oxide catalysts convert ethane in LODh processes, they usually demonstrate low activity yet still may cause overoxidation of fed ethane to COX, thereby leading to low C2H4 yield. Without being bound by any particular theory, it is believed that low conversion may be a result of low catalyst stability caused by the chemical changes, reduction, and volatilization of elements such as tellurium, which is generally encountered when these catalysts are operated in a cyclic redox mode or in conditions with low partial pressure of O2 typically at the outlet sections of fixed bed reactors.

[0050] Reference will now be made in detail to embodiments of the product stream yielded by the alkane-to alkene conversion utilizing the MoVyOxbased dehydrogenation catalysts disclosed herein.

[0051] According to embodiments, the product stream may comprise at least one alkene. In one or more embodiments, the product stream may comprise greater than or equal to 3 vol.% alkenes, such as 10 vol.% alkenes, 20 vol.% alkenes, or 30 vol.% alkenes. In one or more embodiments, the at least one alkenes may consist essentially ethylene, such that the productstream may comprise greater than or equal to 3 vol.% ethylene, such as 10 vol.% ethylene, 20 vol.% ethylene, or 30 vol.% ethylene.

[0052] The product stream may further comprise at least one of oxygenates, carbon monoxide, carbon dioxide, steam, and wherein the product stream is further processed to remove the at least one of at least one of oxygenates, carbon monoxide, carbon dioxide, steam, and alkanes from the product stream.

[0053] In further embodiments, the product stream may comprise less than or equal to 25 vol.% gaseous COX, such as less than or equal to 18 vol.% or less than or equal to 10 vol.%. The gaseous COXmay be CO, CO2, or a combination thereof.

[0054] In further embodiments, the product stream may be substantially free of an oxygenated molecule or molecules containing at least one oxygenated functional group, such as a carboxyl group, a carbonyl group, a hydroxyl group, a ketone group, and / or an epoxy group, meaning that the product stream comprises less than 1 vol.% an oxygenated molecule such as a carboxyl molecule, a carbonyl molecule, or an alcohol molecule. In one or more embodiments, the carboxyl molecule, the carbonyl molecule, or the alcohol molecule may be acetic acid, acetaldehyde, methanol, ethanol, or other light-weight molecule containing at least one carboxyl (-COOH), carbonyl (-C=O), or hydroxyl (-OH) group.

[0055] Reference will now be made in detail to embodiments of regenerating spent dehydrogenation catalysts and forming regenerated dehydrogenation catalysts.

[0056] In both ODH and LODh processes, spent dehydrogenation catalysts are likely to lose lattice oxygen as a result of the conversion. The lost lattice oxygen can be regenerated by reoxidizing the dehydrogenation catalysts with a gaseous oxidant, such as air or oxygen, in a spatially or temporally separated step. By avoiding co-feeding a gaseous oxidant and highly flammable ethane, safety risks are significantly reduced. Further, by using a gaseous oxidant such as air to regenerate the dehydrogenation catalysts, the need for separate, energy-consuming O2 production can be avoided.

[0057] According to embodiments, methods of converting alkanes to alkenes may further comprise contacting spent dehydrogenation catalyst with a regenerative stream, thereby regenerating the spent dehydrogenation catalyst and forming regenerated dehydrogenation catalyst, wherein the regenerative stream may comprise oxygen. In one or more embodiments, the regenerative stream may comprise from 2 vol.% to 22 vol.% oxygen, such as from 2 vol.% to 10 vol.% oxygen, from 7 vol.% to 12 vol.% oxygen, from 7 vol.% to 17 vol.%, from 7 vol.% to 22 vol.% oxygen, from 12 vol.% to 17 vol.% oxygen, from 12 vol.% to 22 vol.% oxygen, or from 17 vol.% to 22 vol.% oxygen. The regenerative stream may be diluted or undiluted air.

[0058] In one or more embodiments, contacting spent dehydrogenation catalyst with a regenerative stream may comprise introducing the regenerative stream and the spent dehydrogenation catalyst into a regeneration zone at a temperature of from 300 °C to 600 °C and a pressure of from 0 bar(g) (0 kPa) to 21 bar(g) (2100 kPa). For example, the temperature may be of from 300 °C to 500 °C, from 400 °C to 550 °C, from 450 °C to 600 °C, or from 500 °C to 600 °C. For example, the pressure may be of from 0 bar(g) to 6 bar(g), from 0 bar(g) to 11 bar(g), from 0 bar(g) to 17 bar(g), from 6 bar(g) to 17 bar(g), from 11 bar(g) to 17 bar(g), from 11 bar (g) to 21 bar(g), or from 17 bar(g) to 21 bar(g).

[0059] The regenerative stream may, in embodiments, be added to the reactor sequentially to the feed stream, such that alkanes and gaseous oxidants are not present in the reaction zone at the same time. It should be understood that, in embodiments, the regenerative stream may be added at various points within the reaction process. This may be accomplished by introducing the regenerative stream into the reactor at different locations within the reactor and / or introducing the regenerative stream at different time periods while the reaction is taking place.

[0060] In one or more embodiments, the regenerative stream may be contacted with the spent dehydrogenation catalyst at a weight hourly space velocity of 0.1 to 10 per hour.

[0061] According to embodiments, contacting the spent dehydrogenation catalyst with the regenerative stream regenerates the spent dehydrogenation catalysts, thereby forming regenerated dehydrogenation catalysts that can be used in the reaction zone to aid the conversion of alkanes to alkenes.

[0062] In embodiments, upon contacting the spent dehydrogenation catalyst with the regenerative stream it would replenish the structure with oxygen in the process recovering from 0.05 wt.% to 5 wt.% of the spent dehydrogenation catalyst.

[0063] In one or more embodiments, the method of converting alkanes to alkenes may further comprise removing spent dehydrogenation catalyst from the reaction zone and introducing the spent dehydrogenation catalyst into a regeneration zone. In embodiments, the regeneration zone may be part of the reactor. In embodiments, the regeneration zone may be separate from the reactor.

[0064] In embodiments, the method of converting alkanes to alkenes may further comprise regenerating the spent dehydrogenation catalyst, thereby forming regenerated dehydrogenation catalyst. In embodiments, regenerating the spent dehydrogenation catalyst may comprise contacting the spent dehydrogenation catalyst with a regeneration stream comprising gaseous oxygen, air, or combinations thereof. In some embodiments, the regeneration zone is purged with gaseous nitrogen prior to contacting the spent dehydrogenation catalyst with the regeneration stream. In some embodiments, before entering the regeneration zone, the spent catalyst is stripped of the entrained hydrocarbons with inert gas such as nitrogen prior to contacting the spent dehydrogenation catalyst with the regeneration stream. In embodiments, the dehydrogenation catalyst may be regenerated at a temperature of less than or equal to 600 °C, such as from 20 °C to 600 °C, from 25 °C to 600 °C, from 50 °C to 600 °C, from 100 °C to 600 °C, from 250 °C to 600 °C, from 500 °C to 600 °C, from 20 °C to 500 °C, from 25 °C to 500 °C, from 50 °C to 500 °C, from 100 °C to 500 °C, from 250 °C to 500 °C, from 20 °C to 250 °C, from 25 °C to 250 °C, from 50 °C to 250 °C, from 100 °C to 250 °C, from 20 °C to 100 °C, from 25 °C to 100 °C, from 50 °C to 100 °C, from 20 °C to 50 °C, or from 25 °C to 50 °C. In one or more embodiments, the dehydrogenation catalyst may be regenerated for a time of at least 1 second, such as from 1 second to 60 minutes, from 1 second to 45 minutes, from 1 second to 30 minutes, from 1 second to 20 minutes, from 1 second to 10 minutes, from 1 second to 5 minutes, from 1 second to 1 minute, from 1 second to 30 seconds, from 1 second to 15 seconds, from 1 second to 10 seconds, from 1 second to 5 seconds, or from 1 second to two seconds. In embodiments, the dehydrogenation catalyst may be heated prior to sending the dehydrogenation catalyst back to the reaction zone.

[0065] In one or more embodiments, the method of converting alkanes to alkenes may further comprise returning regenerated dehydrogenation catalyst to the reaction zone where it is contacted with the feed stream.EXAMPLES

[0066] In the below examples, the designation “CE” refers to Comparative Examples and the designation “IE” refers to Inventive Examples.

[0067] CE 1 : Mo V0.3Nb0.01Bi0.1Ox

[0068] Weighed quantities of MoOs (3.0110 g), V2O5 (0.5706 g), ammonium niobium oxalate (0.9351 g), and Bi2Os (0.4854 g) were added to a 125 cm3Teflon-insert autoclave (Model 4748 General Purpose Acid Digestion Vessel, Parr). After adding 37.7 cm3of deionized H2O, the slurry was stirred using a Teflon-coated magnetic stirrer at a speed of 5 rotations per second. Thereafter, 0.174 cm3of ethylene glycol was added under stirring. Finally, 0.6009 g of oxalic acid dihydrate and 0.2651 g of citric acid were added to the slurry. After being stirred for at least 360 seconds, the autoclave was sealed and latched on to a rotating shaft oven. For the hydrothermal synthesis, the temperature was increased at 3 °C min'1(3 K / min) to 200 °C (473 K) and kept at this temperature for 48 hours (172800 s). During the hydrothermal synthesis, the autoclaves were tumbled over a horizontal axis at a speed of 10 rotations per minute (0.167 rotations per second). The obtained material was filtered over a 0.45 pm filter paper and washed with at least 250 cm3of deionized water by pouring over the retentate. The retentate was dried overnight in an oven maintained at 80 °C (353 K).

[0069] The dried sample was heat-treated in a sealed metal bucket inside an oven with a constant flow of N2 (>16 cm3 / s measured at ambient conditions). The heat-treatment procedure consisted of ramping the temperature to 450 °C (723 K) at 5 °C min'1(5 K / min) and holding the temperature at 450 °C (723 K) for 2 hours (7200 s) before cooldown. A similar procedure was followed for three other batches. All four batches were physically homogenized and used for further treatment.

[0070] The heat-treated material was compacted under 7 tonne pressure and sieved to 177 - 400 pm (40-80 mesh) prior to loading in the reactor. The fines were analyzed with X-ray diffraction as described hereinbelow.

[0071] CE2: MoV0.2Nb0.005Sb0.05Bi0.11Ox

[0072] Weighed quantities of MoOs (3.0110 g), V2O5 (0.4661 g), ammonium niobium oxalate (0.0573 g), Sb2O3(0.1858 g), and Bi2O3 (0.4854 g) were added to a 125 cm3Teflon-insert autoclave (Model 4748 General Purpose Acid Digestion Vessel, Parr). After adding 46.2 cm3of deionized H2O, the slurry was stirred using a Teflon-coated magnetic stirrer at a speed of 5 rotations per second. Thereafter, 0.106 cm3of ethylene glycol was added under stirring. Finally, 0.6494 g of oxalic acid dihydrate and 0.3692 g of citric acid were added to the slurry. After being stirred for at least 360 seconds, the autoclave was sealed and latched on to a rotating shaft oven. For the hydrothermal synthesis, the temperature was increased at 3 °C min'1(3 K / min) to 200 °C (473 K) and kept at this temperature for 48 hours (172800 s). During the hydrothermal synthesis, the autoclaves were tumbled over a horizontal axis at a speed of 10 rotations per minute (0.167 rotations per second). The obtained material was filtered over a 0.45 pm filter paper and washed with at least 250 cm3of deionized water by pouring over the retentate. The retentate was dried overnight in an oven maintained at 80 °C.

[0073] The dried sample was heat-treated in a sealed metal bucket inside an oven with a constant flow of N2(>16 cm3 / s measured at ambient conditions). The heat-treatment procedure consisted of ramping the temperature to 450 °C (723 K) at 5 °C min'1(5 K / min) and holding the temperature at 450 °C (723 K) for 2 hours (7200 s) before cooldown.

[0074] The heat-treated material was compacted under 7 tonne pressure and sieved to 177 - 400 pm (40-80 mesh) prior to loading in the reactor. The fines were analyzed with X-ray diffraction as described hereinbelow.

[0075] CE3: MoVo.iNbo.oosSbo.osTbo.oosOx

[0076] Weighed quantities of MoOs (3.6891 g), V2Os (0.4706 g), ammonium niobium oxalate (0.0577 g), Sb2O3 (0.1877 g), and Tb4O? (0.0241 g) were added to a 125 cm3Teflon-insertautoclave (Model 4748 General Purpose Acid Digestion Vessel, Parr). After adding 46 cm3of deionized H2O, the slurry was stirred using a Teflon-coated magnetic stirrer at a speed of 5 rotations per second. Thereafter, 0.107 cm3of ethylene glycol was added under stirring. Finally, 0.6491 g of oxalic acid dihydrate and 0.3692 g of citric acid were added to the slurry. After being stirred for at least 360 seconds, the autoclave was sealed and latched on to a rotating shaft oven. For the hydrothermal synthesis, the temperature was increased at 3 °C min'1(3 K / min) to 200 °C (473 K) and kept at this temperature for 48 hours (172800 s). During the hydrothermal synthesis, the autoclaves were tumbled over a horizontal axis at a speed of 10 rotations per minute (0.167 rotations per second). The obtained material was filtered over a 0.45 pm filter paper and washed with at least 250 cm3of deionized water by pouring over the retentate. The retentate was dried overnight in an oven maintained at 80 °C (353 K).

[0077] The dried sample was heat-treated in a sealed metal bucket inside an oven with a constant flow of N2 (>16 cm3 / s measured at ambient conditions). The heat-treatment procedure consisted of ramping the temperature to 450 °C (723 K) at 5 °C min'1(5 K / min) and holding the temperature at 450 °C (723 K) for 2 hours (7200 s) before cooldown.

[0078] The heat-treated material was compacted under 7 tonne pressure and sieved to 177 - 400 pm (40-80 mesh) prior to loading in the reactor. The fines were analyzed with X-ray diffraction as described hereinbelow.

[0079] CE4: MoVo.iNbo.oosSbo.oiTbo. Ox

[0080] Weighed quantities of MoOs (3.6890 g), V2O5 (0.4706 g), ammonium niobium oxalate (0.0576 g), Sb2O3(0.0372 g), and Tb4O? (0.6747 g) were added to a 125 cm3Teflon-insert autoclave (Model 4748 General Purpose Acid Digestion Vessel, Parr). After adding 46 cm3of deionized H2O, the slurry was stirred using a Teflon-coated magnetic stirrer at a speed of 5 rotations per second. Thereafter, 0.107 cm3of ethylene glycol was added under stirring. Finally, 0.6489 g of oxalic acid dihydrate and 0.3693 g of citric acid were added to the slurry. After being stirred for at least 360 seconds, the autoclave was sealed and latched on to a rotating shaft oven. For the hydrothermal synthesis, the temperature was increased at 3 °C min'1(3 K / min) to 200 °C (473 K) and kept at this temperature for 48 hours (172800 s). During the hydrothermal synthesis, the autoclaves were tumbled over a horizontal axis at a speed of 10 rotations per minute (0.167rotations per second). The obtained material was filtered over a 0.45 pm filter paper and washed with at least 250 cm3of deionized water by pouring over the retentate. The retentate was dried overnight in an oven maintained at 80 °C (353 K).

[0081] The dried sample was heat-treated in a sealed metal bucket inside an oven with a constant flow of N2 (>16 cm3 / s measured at ambient conditions). The heat-treatment procedure consisted of ramping the temperature to 450 °C (723 K) at 5 °C min'1(5 K / min) and holding the temperature at 450 °C (723 K) for 2 hours (7200 s) before cooldown.

[0082] The heat-treated material was compacted under 7 tonne pressure and sieved to 177 - 400 pm (40-80 mesh) prior to loading in the reactor. The fines were analyzed with X-ray diffraction as described hereinbelow.

[0083] IE1: MoV0.20Nb0.005Sb0.05Tb0.10Ox

[0084] Weighed quantities of MoOs (1.4824 g), V2O5 (0.1891 g), ammonium niobium oxalate (0.0233 g), Sb2O3(0.0751 g), and Tb4O? (0.1935 g) were added to a45 cm3Teflon-insert autoclave (Model 4744 General Purpose Acid Digestion Vessel, Parr). After adding 18.6 cm3of deionized H2O, the slurry was stirred using a Teflon-coated magnetic stirrer at a speed of 13.3 rotations per second. Thereafter, 0.043 cm3of ethylene glycol was added under stirring. Finally, 0.2609 g of oxalic acid dihydrate and 0.1484 g of citric acid were added to the slurry. After being stirred for at least 360 seconds, the autoclave was sealed and latched on to a rotating shaft oven. For the hydrothermal synthesis, the temperature was increased at 3 °C min'1(3 K / min) to 200 °C (473 K) and kept at this temperature for 48 hours (172800 s). During the hydrothermal synthesis, the autoclaves were tumbled over a horizontal axis at a speed of 10 rotations per minute (0.167 rotations per second). The obtained material was filtered over a 0.45 pm filter paper and washed with at least 90 cm3of deionized water by pouring over the retentate. The retentate was dried overnight in an oven maintained at 80 °C (353 K).

[0085] The dried sample was heat-treated in a sealed metal bucket inside an oven with a constant flow of N2 (>16 cm3 / s measured at ambient conditions). The heat-treatment procedure consisted of ramping the temperature to 450 °C (723 K) at 5 °C min'1(5 K / min) and holding the temperature at 450 °C (723 K) for 2 hours (7200 seconds) before cooldown.

[0086] The heat-treated material was compacted under 7 tonne pressure and sieved to 177 - 400 pm (40-80 mesh) prior to loading in the reactor. The fines were analyzed with X-ray diffraction as described hereinbelow.

[0087] IE2: MoV0.2Ta0.005Sb0.05Tb0.10Ox

[0088] Weighed quantities of MoO3(3.6889 g), V2O5 (0.4707 g), Ta2O5(0.0286 g), Sb2O3(0.1876 g), and Tb4O? (0.4817 g) were added to a 125 cm3Teflon-insert autoclave (Model 4748 General Purpose Acid Digestion Vessel, Parr). After adding 46 cm3of deionized H2O, the slurry was stirred using a Teflon-coated magnetic stirrer at a speed of 5 rotations per second. Thereafter, 0.107 cm3of ethylene glycol was added under stirring. Finally, 0.6493 g of oxalic acid dihydrate and 0.3689 g of citric acid were added to the slurry. After being stirred for at least 360 seconds, the autoclave was sealed and latched on to a rotating shaft oven. For the hydrothermal synthesis, the temperature was increased at 3 °C min'1(3 K / min) to 200 °C (473 K) and kept at this temperature for 48 hours (172800 s). During the hydrothermal synthesis, the autoclaves were tumbled over a horizontal axis at a speed of 10 rotations per minute (0.167 rotations per second). The obtained material was filtered over a 0.45 pm filter paper and washed with at least 250 cm3of deionized water by pouring over the retentate. The retentate was dried overnight in an oven maintained at 80 °C (353 K).

[0089] The dried sample was heat-treated in a sealed metal bucket inside an oven with a constant flow of N2(>16 cm3 / s measured at ambient conditions). The heat-treatment procedure consisted of ramping the temperature to 450 °C (723 K) at 5 °C min'1(5 K / min) and holding the temperature at 450 °C (723 K) for 2 hours (7200 s) before cooldown.

[0090] The heat-treated material was compacted under 7 tonne pressure and sieved to 177 - 400 pm (40-80 mesh) prior to loading in the reactor. The fines were analyzed with X-ray diffraction as described hereinbelow.

[0091] IE3: MoV0.2Nb0.005Sb0.05Tb0.20Ox

[0092] Weighed quantities of MoOs (3.4003 g), V2Os (0.4337 g), ammonium niobium oxalate (0.0533 g), Sb2O3 (0.1730 g), and Tb4O? (0.8881 g) were added to a 125 cm3Teflon-insertautoclave (Model 4748 General Purpose Acid Digestion Vessel, Parr). After adding 42.4 cm3of deionized H2O, the slurry was stirred using a Teflon-coated magnetic stirrer at a speed of 5 rotations per second. Thereafter, 0.099 cm3of ethylene glycol was added under stirring. Finally, 0.5979 g of oxalic acid dihydrate and 0.3402 g of citric acid were added to the slurry. After being stirred for at least 360 seconds, the autoclave was sealed and latched on to a rotating shaft oven. For the hydrothermal synthesis, the temperature was increased at 3 °C min'1(3 K / min) to 200 °C (473 K) and kept at this temperature for 48 hours (172800 s). During the hydrothermal synthesis, the autoclaves were tumbled over a horizontal axis at a speed of 10 rotations per minute (0.167 rotations per second). The obtained material was filtered over a 0.45 pm filter paper and washed with at least 250 cm3of deionized water by pouring over the retentate. The retentate was dried overnight in an oven maintained at 80 °C (353 K).

[0093] The dried sample was heat-treated in a sealed metal bucket inside an oven with a constant flow of N2 (>16 cm3 / s measured at ambient conditions). The heat-treatment procedure consisted of ramping the temperature to 450 °C (723 K) at 5 °C min'1(5 K / min) and holding the temperature at 450 °C (723 K) for 2 hours (7200 s) before cooldown.

[0094] The heat-treated material was compacted under 7 tonne pressure and sieved to 177 - 400 pm (40-80 mesh) prior to loading in the reactor. The fines were analyzed with X-ray diffraction as described hereinbelow.

[0095] IE4: MoV0.2Nb0.005Sb0.10Tb0.10Ox

[0096] Weighed quantities of MoOs (3.6887 g), V2O5 (0.4707 g), ammonium niobium oxalate (0.0577 g), Sb2O3(0.3752 g), and Tb4O? (0.4817 g) were added to a 125 cm3Teflon-insert autoclave (Model 4748 General Purpose Acid Digestion Vessel, Parr). After adding 46 cm3of deionized H2O, the slurry was stirred using a Teflon-coated magnetic stirrer at a speed of 5 rotations per second. Thereafter, 0.107 cm3of ethylene glycol was added under stirring. Finally, 0.6492 g of oxalic acid dihydrate and 0.3689 g of citric acid were added to the slurry. After being stirred for at least 360 seconds, the autoclave was sealed and latched on to a rotating shaft oven. For the hydrothermal synthesis, the temperature was increased at 3 °C min'1(3 K / min) to 200 °C (473 K) and kept at this temperature for 48 hours (172800 s). During the hydrothermal synthesis, the autoclaves were tumbled over a horizontal axis at a speed of 10 rotations per minute (0.167rotations per second). The obtained material was filtered over a 0.45 pm filter paper and washed with at least 250 cm3of deionized water by pouring over the retentate. The retentate was dried overnight in an oven maintained at 80 °C (353 K).

[0097] The dried sample was heat-treated in a sealed metal bucket inside an oven with a constant flow of N2 (>16 cm3 / s measured at ambient conditions). The heat-treatment procedure consisted of ramping the temperature to 450 °C (723 K) at 5 °C min'1(5 K / min) and holding the temperature at 450 °C (723 K) for 2 hours (7200 s) before cooldown.

[0098] The heat-treated material was compacted under 7 tonne pressure and sieved to 177 - 400 pm (40-80 mesh) prior to loading in the reactor. The fines were analyzed with X-ray diffraction as described hereinbelow.

[0099] Table 1 shows amounts of additives added to prepare the catalysts of the Examples.Table 1: Additives to Prepare the Catalysts

[0100] MEASUREMENT INFORMATION

[0101] The C2H4 production rate (PC2H4) at any given point of time is defined by Equation (1), wherein mcataiyst is the mass of the loaded catalyst in grams (0.350 g in the examples hereinabove),FOut,A is the inlet molar flow rate of component ‘A’ in mol s’1, and MMA is the molar mass of component ‘A’ in g mol’1:

[0102] The total fraction of oxygen removed (O) at any given point of time (say, ti) is defined by Equation (2), while taking an integral from zero to ti, wherein mcataiyst is the mass of the loaded catalyst in grams (0.350 g in the examples hereinabove), F0Ut,A is the inlet molar flow rate of component ‘A’ in mol s’1, and MMA is the molar mass of component ‘A’ in g mol’1:Equation (2)

[0103] The above equation assumes the following reactions:C2H6+ [0] C2H4+ H2O Reaction (1)

[0104] Partial oxidation of C2H6 to CO and H2 was excluded in the calculations because negligible quantities of H2 were observed during testing. [O] in the above equations represents lattice oxygen from the catalyst.

[0105] The catalyst to ethane ratio (CE) is defined by Equation (3), wherein mcataiyst is the mass of the loaded catalyst in grams (0.350 g), TOS is the time on stream in seconds, Fin,c2H6is the inlet molar flow rate in mol / s, and MMC2H6is the molar mass of C2H6 (30 g / mol):

[0106] The cumulative ethane conversion (XC2H6) was calculated using Equation (4), wherein TOS is between 10 s and 40 s, corresponding to CE between 56.5 and 226, nin,c2H6is the total amount of ethane fed during this duration, and n0Ut,c2H6is the total amount of ethane detected in the reactor effluent: 100 Equation (4)

[0107] The cumulative C2H4 selectivity (SC2H4) was calculated using Equation (5), wherein TOS is between 10 s and 40 s, corresponding to CE between 56.5 and 226, nin,c2H6is the total amount of ethane fed during this duration, n0Ut,c2H4is the total amount of C2H4 produced, also known as cumulative amount of C2H4 produced:

[0108] Selectivity to CO (Sco) and CO2 (Sco2) are calculated on a similar carbon-based calculation.

[0109] The cumulative C2H4 yield (YC2H4) was calculated using the product of conversion and selectivity according to Equation (6):

[0110] TEST METHODS

[0111] CATALYST PERFORMANCE TESTING

[0112] Performance testing was performed in a fixed bed high-throughput reactor set-up from Sintef (High Pressure Reactor Assembly Module), with parallel SS316 reactor tubes (I.D. 3 mm) located in a heated block. The reactor system was equipped with an online GC Siemens Maxum- II instrument. For catalytic testing, 350 mg of catalyst particles were loaded in the reactor. The total pressure was kept constant at 250 kPa (2.5 bar) with a back pressure regulator and thetemperature was kept constant at 450 °C (723 K). In the LODh mode of operation, 50% C2H6 in N2 / He was fed for 360 seconds (6 minutes) at 0.25 cm3s'1or 15 mL min'1(measured at ambient conditions) at a weight hourly space velocity close to 1.6 hr’1. Thereafter, N2 was fed to purge the reactors for 600 seconds (10 minutes) at 0.167 cm3 / s or 10 mi min'1(measured at ambient conditions). The catalysts were then regenerated by feeding 2.5% O2 in N2 for 4500 seconds (75 minutes) at 0.167 cm3 / s or 10 ml min'1(measured at ambient conditions). After purging the reactor again with N2, a new cycle began.

[0113] The zero time or the time at which it is assumed that ethane contacts the catalyst was estimated using a thermal conductivity detector (TCD). During this estimation, the entire reactor effluent stream was directed to the TCD, which is highly sensitive to gas composition. The zero time was estimated by switching the flow from N2 to C2H6 at 353 K (80 °C) and 250 kPa (2.5 bar). The point at which the TCD signal reaches 75% of its asymptotic values is assumed to be the zero time. The measurement was repeated 10 times and the average zero time was used for further calculations. The results were as shown in Table 2 and Table 3, as well as in FIGS. 1-4.Table 2: C2H4 Production Rate with Depletion of OxygenTable 3: Conversion, Selectivity, and Yield

[0114] X-RAY DIFFRACTION (XRD) MEASUREMENTS

[0115] XRD of finely powdered catalysts (approximately 100 mg, less than 177 pm) were obtained using a Bruker D8 Discover diffractometer working at 40 mA and 40 kV using a graphite monochromator and Cu Ka 1,2 source (A = 1.5418 A) and Vantec 500 General Area Diffraction Detector System (GADDS) detector at ambient temperature and pressure. The diffractograms were recorded for 29 angles ranging from 2 to 80° with a step size of 0.0195° and 0.5 s integration step size. The results of XRD for CE2 and IE1 were as shown in FIG. 5.

[0116] A comparison of IE1-IE4 with CE1-CE4 in Table 2 shows that materials without Tb (CE1, CE2), with very little terbium (CE3), and with terbium but very little antimony (CE4) have lower C2H4 productivity at higher levels of oxygen depletion (>0.7 mass%). Higher oxygen transport capacity along with higher C2H4 productivity are critical to catalyst performance and process economics.

[0117] Table 3 and FIG. 2 show that during the initial period of low oxygen depletion, CE2 performs better than IE1. However, the production rates of C2H4 remain high for IE1 when oxygen depletion is higher. The latter corresponds to longer time duration or time on stream (see FIG. 3). To estimate cumulative C2H4 production over a longer duration, Table 3 provides evidence that IE1 produces at least as much C2H4 over a longer duration as CE2. Table 2 and FIG. 3 show that, for IE1-IE4, the productivity of C2H4 and yield of C2H4 remain consistently high over the duration when experimental measurements were possible. There is no indication that the yield of C2H4 decreases with time for IE1-IE4, unlike CE1-CE4. Thus, it is reasonable to expect that over an even longer duration, the total amount of C2H4 produced with IE1 will exceed that in CE2.

[0118] CE3 contains very little terbium (atomic Mo / Tb = 0.005) and this sample shows very high ethane conversion with low selectivity to C2H4. Over time, this material tends to sinter in the reactor causing an increased pressure drop over the catalyst bed, which makes experimental measurements unreliable. CE4 contains terbium (Mo / Tb = 0.14) but very little antimony (Mo / Sb = 0.01). Interestingly, this causes a steep drop in C2H4 with increasing oxygen depletion (Table 2 and FIG. 1). This suggests that sufficient amounts of both terbium and antimony may help to sustain C2H4 production at higher levels of oxygen depletion. IE3 with the highest proportion of terbium (Mo / Tb = 0.20) shows the highest selectivity to C2H4 (around 82%), which remainsconsistent even at higher oxygen depletion. Low selectivity to COXin IE3 is advantageous for limiting the overall CO2 emissions of the process using the discussed catalyst.

[0119] IE2 has a composition similar to IE1 but has niobium substituted for tantalum. This substitution resulted in decreased activity for ethane conversion with little to no improvement in selectivity to C2H4, thereby leading to lower C2H4 yield. Excluding IE2 and plotting C2H4 yield over a limited initial period vs the nominal atomic ratio of terbium to antimony shows that an optimal terbium to antimony atomic ratio likely exists within a range of 0.1 and 14 (FIG. 4).

[0120] A surprising finding for the terbium samples in conjunction with antimony are their capability to sustain C2H4 production rates, which, without being bound by theory, is believed to be caused by faster diffusion of lattice oxygen to the surface of a dehydrogenation catalyst.

[0121] The subject matter of the present disclosure has been described in detail and by reference to specific embodiments. It should be understood that any detailed description of a component or feature of an embodiment does not necessarily imply that the component or feature is essential to the particular embodiment or to any other embodiment. Further, it should be apparent to those skilled in the art that various modifications and variations can be made to the described embodiments without departing from the spirit and scope of the claimed subject matter.

[0122] It is noted that one or more of the following claims utilize the term “wherein” as a transitional phrase. For the purposes of defining the present technology, it is noted that this term is introduced in the claims as an open-ended transitional phrase that is used to introduce a recitation of a series of characteristics of the structure and should be interpreted in like manner as the more commonly used open-ended preamble term “comprising.”

[0123] It should be understood that where a first component is described as “comprising” a second component, it is contemplated that, in embodiments, the first component “consists” or “consists essentially of’ that second component. It should further be understood that where a first component is described as “comprising” a second component, it is contemplated that, in embodiments, the first component comprises at least 10%, at least 20%, at least 30%, at least 40%, at least 50%, at least 60%, at least 70%, at least 80%, at least 90%, at least 95%, or even at least 99% that second component (where % can be weight % or molar %).

[0124] It is also noted that recitations herein of “at least one” component, element, etc., should not be used to create an inference that the alternative use of the articles “a” or “an” should be limited to a single component, element, etc.

Claims

CLAIMS1. A method for converting alkanes to alkenes, the method comprising: contacting a feed stream comprising alkanes with a dehydrogenation catalyst in a reaction zone, the dehydrogenation catalyst comprising molybdenum, vanadium, at least 0.1 wt.% terbium, and at least 0.5 wt.% antimony, wherein the weight percent is based on a total weight of the dehydrogenation catalyst; and converting at least a portion of the alkanes to alkenes, thereby yielding a product stream comprising alkanes, alkenes, and steam.

2. The method of claim 1, wherein an atomic ratio of terbium to antimony is from 0.05: 1 to 20:1.

3. The method of any of the preceding claims, wherein: the feed stream and / or the reactor are substantially free of a gaseous oxidant; or the feed stream and / or the reactor comprise oxygen.

4. The method of any of the preceding claims, wherein the dehydrogenation catalyst is substantially free of tellurium (Te).

5. The method of any one of the preceding claims, wherein the dehydrogenation catalyst further comprises niobium, tantalum, or combinations thereof.

6. The method of claim 5, wherein: the dehydrogenation catalyst comprises at least 0.04 wt.% of niobium, tantalum, or combinations thereof.

7. The method of claim 5 or claim 6, wherein the dehydrogenation catalyst comprises: an atomic ratio of niobium to molybdenum of from 0: 1 to 0.1 : 1; an atomic ratio of tantalum to molybdenum of from 0: 1 to 0.1 : 1; and a sum of the atomic ratios of niobium to molybdenum and tantalum to molybdenum of from 0.001 : 1 to 0.1 : 1.

8. The method of any one of the preceding claims, wherein the dehydrogenation catalyst comprises a crystal structure with a Pba2-32 space group.9 The method of any one of the preceding claims, wherein the dehydrogenation catalyst comprises:33 wt.% to 66 wt.% molybdenum;2 wt.% to 16 wt.% vanadium;0.1 wt.% to 19 wt.% terbium; and0.5 wt.% to 14 wt.% antimony; wherein the weight percent is based on a total weight of the dehydrogenation catalyst.

10. The method of claim 9, wherein the dehydrogenation catalyst comprises: an atomic ratio of vanadium to molybdenum of from 0.1 : 1 to 0.6: 1 ; an atomic ratio of terbium to molybdenum of from 0.002: 1 to 0.2: 1; and an atomic ratio of antimony to molybdenum of from 0.01 : 1 to 0.2: 1.

11. The method of any one of the preceding claims, wherein the molybdenum, vanadium, and terbium are present in the dehydrogenation catalyst as oxides.

12. The method of any one of the preceding claims, wherein the converting at least a portion of the alkanes to alkenes occurs at a pressure from 0 bar(g) to 20 bar(g) at a WHSV of from 0.1 h'1to 10 h’1, and at a temperature that is less than or equal to 600 °C.

13. The method of any one of the preceding claims, wherein the dehydrogenation catalyst and the alkanes in the feed stream have a mass to mass ratio that is from 5 : 1 to 500: 1.

14. The method of any one of the preceding claims, wherein the method further comprises: contacting spent dehydrogenation catalyst with a regenerative stream, thereby regenerating the spent dehydrogenation catalyst and forming regenerated dehydrogenation catalyst, wherein the regenerative stream comprises oxygen.

15. The method of claim 14, wherein contacting the spent dehydrogenation catalyst with the regenerative stream comprises introducing the regenerative stream and the spent dehydrogenation catalyst into a regeneration zone at a temperature of from 300 °C to 600 °C and a pressure of from 0 bar(g) (0 kPa) to 21 bar(g) (2100 kPa).

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