Limiting acetic acid production in ethane oDH processes

Abstract

The present disclosure relates to limiting the production of acetic acid in an oxidative dehydrogenation process of converting ethane to ethylene. The process of oxidative dehydrogenation includes feeding acetic acid along with ethane and oxygen into an oxidative dehydrogenation reactor, where contact with a catalyst results in the conversion of ethane to ethylene and acetic acid. By including acetic acid in the feed, the amount of acetic acid produced can be limited, and the ratio of ethylene produced to ethane consumed can be improved.

Description

[0001] Priority requirements This application claims priority to U.S. Provisional Application No. 63 / 036,471, filed June 9, 2020, the entire contents of which are incorporated herein by reference. Technical Field

[0002] This specification relates to an oxidative dehydrogenation method for the conversion of ethane to ethylene. More specifically, an oxidative dehydrogenation method is described in which acetic acid is included in the feed stream. Background Technology

[0003] Oxidative dehydrogenation of ethane using mixed metal oxide catalysts is an alternative to steam cracking for ethylene production. Oxidative dehydrogenation of ethane is a method for converting relatively inert ethane into the more reactive and valuable ethylene. It involves the endothermic removal of hydrogen from ethane and the exothermic oxidation of hydrogen to produce water. However, oxidative dehydrogenation of ethane can also lead to the production of unwanted byproducts such as acetic acid. Therefore, selective oxidation of hydrogen and minimization of ethane oxidation can advantageously maximize the production of the desired product, ethylene. Summary of the Invention

[0004] Certain aspects of the described subject matter can be implemented as a method for oxidative dehydrogenation (ODH) of ethane. The method involves contacting a feed stream containing ethane, oxygen, and acetic acid with an oxidative dehydrogenation catalyst in an oxidative dehydrogenation reactor under oxidative dehydrogenation conditions to produce a product stream containing ethylene, unreacted ethane, water, and acetic acid. The concentration of acetic acid in the feed stream is 0.5 to 10% by volume. The feed stream may optionally contain an inert diluent. In some cases, the product stream contains carbon monoxide, carbon dioxide, or both.

[0005] This aspect and other aspects may include one or more of the following features.

[0006] The concentration of acetic acid in the feed stream can be 2 to 5% by volume. The concentration of acetic acid in the feed stream can be greater than 2% by volume. The feed stream to the reactor can contain 0.5 to 10% by volume of acetic acid. The feed stream to the reactor can contain 0.5 to 0.7 molar ratios of oxygen and ethane. The feed stream to the reactor can contain a certain molar ratio of water (H₂O) and carbon dioxide (CO₂) to ensure that the feed composition is outside the flammability limits.

[0007] The method may include downstream separation methods. For example, the product stream can be separated into a liquid stream and a gaseous component stream by a downstream separation method. The liquid stream may contain water and acetic acid. The gaseous component stream may contain ethylene and unreacted ethane. In some cases, the gaseous component stream contains carbon monoxide, carbon dioxide, or both. At least a portion of the liquid stream can be recycled back to the reactor as part of the feed stream. The liquid stream can be diluted with water to achieve a desired amount of acetic acid in the feed stream. The fractionation rate of the liquid stream can be adjusted to achieve a desired amount of acetic acid in the feed stream.

[0008] The reactor can operate at temperatures ranging from 300°C to 425°C. The reactor can operate at temperatures ranging from 315°C to 400°C. The reactor can operate at pressures ranging from 0.5 psig to 100 psig. The reactor can operate at pressures ranging from 15 psig to 50 psig.

[0009] The gas hourly space velocity (GHSV) of the product stream can be 500 h⁻¹. -1 Up to 30000h -1 The GHSV of the product stream can be 1000h. -1 Up to 150,000h -1 The GHSV of the product stream can be 500h. -1 Up to 4000h -1 .

[0010] The catalyst may include one or more catalysts selected from the following: i) Catalysts of the following formula: Mo a V b Te c Nb d Pd e O f Where: a, b, c, d, e, and f are the relative atomic weights of elements Mo, V, Te, Nb, Pd, and O, respectively; and when a = 1, b = 0.01 to 1.0, c = 0.01 to 1.0, d = 0.01 to 1.0, and 0.00 ≤ e ≤ 0.10, f is the number of valence states that satisfy the catalyst. ii) Catalysts of the following formula: Mo a E k G l O f Where: E is selected from Ba, Ca, Cr, Mn, Nb, Ta, Ti, Te, V, W and their mixtures; G is selected from Bi, Ce, Co, Cu, Fe, K, Mg, V, Ni, P, Pb, Sb, Si, Sn, Ti, U and their mixtures; a=1; k is 0 to 2; l=0 to 2, provided that the total value of l for CO, Ni, Fe and their mixtures is less than 0.5; and f is the number of catalyst valence states. iii) Catalysts of the following formula: V m Mo n Nb o Te p Me q O f Where: Me is a metal selected from Ta, Ti, W, Hf, Zr, Sb and mixtures thereof; m is 0.1 to 3; n is 0.5 to 1.5; o is 0.001 to 3; p is 0.001 to 5; q is 0 to 2; and f is the number of catalyst valence states; and iv) Catalysts of the following formula: Mo a V r X s Y t Z u M v O f Wherein: X is at least one of Nb and Ta; Y is at least one of Sb and Ni; Z is at least one of Te, Ga, Pd, W, Bi and Al; M is at least one of Fe, Co, Cu, Cr, Ti, Ce, Zr, Mn, Pb, Mg, Sn, Pt, Si, La, K, Ag and In; a = 1.0 (normalized); r = 0.05 to 1.0; s = 0.001 to 1.0; t = 0.001 to 1.0; u = 0.001 to 0.5; v = 0.001 to 0.3; and f is the number of catalyst valence states.

[0011] The liquid stream may contain less than 10% by volume of acetic acid.

[0012] The selectivity for ethylene production can be 75% to 99%. The selectivity for CO2 production can be equal to or less than 10%. The selectivity for CO production can be equal to or less than 11%. The feed stream to the reactor can contain 2 to 3 vol% acetic acid, 29 to 57 vol% H2O, 16 to 26 vol% C2H6, 8 to 14 vol% O2, and 17 to 28 vol% CO2.

[0013] The makeup stream and gaseous component stream can be contacted with the second oxidative dehydrogenation catalyst in the second oxidative dehydrogenation reactor under oxidative dehydrogenation conditions to produce a second product stream. The second product stream may contain ethylene, unreacted ethane, water, and acetic acid. In some cases, the second product stream may contain carbon monoxide, carbon dioxide, or both. The total concentration of acetic acid in the makeup stream and gaseous component stream (together) entering the second oxidative dehydrogenation reactor can be from 0.5% to 10% by volume.

[0014] The method may include a second downstream separation method. For example, the second product stream may be separated into a second liquid stream and a second gaseous component stream by the second downstream separation method. The second liquid stream may contain water and acetic acid. The second gaseous component stream may contain ethylene and unreacted ethane. In some cases, the second gaseous component stream contains carbon monoxide, carbon dioxide, or both. At least a portion of the second liquid stream may be recycled back to the reactor as part of the feed stream. The second liquid stream may be diluted with water to achieve a desired amount of acetic acid in the feed stream. The fractional flow rate of the second liquid stream may be adjusted to achieve a desired amount of acetic acid in the feed stream. At least a portion of the second liquid stream may be recycled back to the second reactor as part of the makeup stream. The second liquid stream may be diluted with water to achieve a desired amount of acetic acid in the makeup stream. The fractional flow rate of the second liquid stream may be adjusted to achieve a desired amount of acetic acid in the makeup stream.

[0015] It should be understood that this disclosure as described in this specification is not limited to the examples summarized in this overview. Various other aspects are described and illustrated herein. Attached Figure Description

[0016] Figure 1 This is a schematic diagram of an example system including an oxidative dehydrogenation reactor.

[0017] Figure 2A For including Figure 1 Schematic diagrams of two implemented examples of oxidative dehydrogenation reactors.

[0018] Figure 2B For including Figure 1 Schematic diagrams of two implemented examples of oxidative dehydrogenation reactors.

[0019] Figure 3A This is a flowchart of an example method for oxidative dehydrogenation.

[0020] Figure 3B This is a flowchart of an example method for oxidative dehydrogenation. Detailed Implementation

[0021] Certain exemplary aspects of this disclosure will now be described to provide a comprehensive understanding of the principles of the methods disclosed herein. It will be understood by those skilled in the art that the methods described herein are non-limiting exemplary aspects, and that the scope of each instance of this disclosure is defined only by the claims. Features described or illustrated in connection with one exemplary aspect may be combined with features of other aspects. These modifications and variations are intended to be included within the scope of this disclosure.

[0022] Except as indicated in the operational embodiments, or otherwise specified, all figures or expressions relating to the amounts of ingredients, reaction conditions, etc., used in the specification and claims should be understood to be modified by the term "about" in all cases. Therefore, unless otherwise specified, the numerical parameters listed in the following specification and appended claims are approximate values, which may vary according to the properties desired by this disclosure. At least, and without attempting to limit the application of the equivalence doctrine to the scope of the claims, each numerical parameter should be interpreted at least according to the reported significant figures and by applying ordinary rounding methods.

[0023] definition As used herein, the term "oxidative dehydrogenation catalyst" refers to a catalyst used in the oxidative dehydrogenation of ethane to ethylene. The most commonly described oxidative dehydrogenation catalyst is a mixed metal oxide catalyst. Unless otherwise specified, the term "catalyst" is used synonymously with oxidative dehydrogenation catalyst. Furthermore, references to oxidative dehydrogenation catalysts may include mixtures of more than one oxidative dehydrogenation catalyst, each having a different chemical composition, whether supported or unsupported.

[0024] As used herein, the term "feed stream" refers to the gas stream that initially contacts the oxidative dehydrogenation catalyst. The feed stream in a typical oxidative dehydrogenation process contains ethane and oxygen components and may contain one or more inert diluents. In some cases, the contribution of a component to the feed stream is described as "feed composition," stated herein as a volume percentage (%) of one or more components. In other cases, volume percentage ratios are used to describe the components within the feed stream, particularly oxygen and ethane.

[0025] As used herein, the term "inert diluent" refers to the gaseous composition used to dilute ethane and oxygen. The inert diluent should be primarily in a gaseous state under oxidative dehydrogenation conditions and should not increase the flammability of ethane. Common inert diluents for oxidative dehydrogenation known to those skilled in the art include, but are not limited to, nitrogen, carbon dioxide, and vapors and mixtures thereof.

[0026] As used herein, the term "under oxidative dehydrogenation conditions" refers to method conditions under which ethane is converted to ethylene by contact with an oxidative dehydrogenation catalyst in the presence of oxygen, including but not limited to temperature, pressure, and feed flow rate. Oxidative dehydrogenation conditions can be adjusted by those skilled in the art to attempt to optimize the conditions for a particular catalyst or whether an inert diluent is used in the feed flow.

[0027] Unless otherwise specified, the term "selectivity" as used herein refers to the selectivity of carbon atoms based on the extent to which ethane is consumed. Selectivity expressed as a percentage can be calculated using the following formula: Where X is the product being evaluated, net mass flow rate refers to the flow rate of X or converted C2H6 in g / min, and is equal to the mass flow rate of X or converted C2H6 in the product stream minus the mass flow rate of component X or converted C2H6 in the feed stream, and molar equivalent (Mol. Equiv.) refers to the amount of X in moles produced completely from 1 mole of ethane or from 1 mole of ethane. If the sum of all selectivities of the products derived from ethane conversion is not 100%, the selectivity is normalized to 100%. Normalization for each product can be calculated by dividing the selectivity of that product by the sum of the selectivities of all carbon atom products.

[0028] The terms “about” or “approximately” as used herein may allow for a degree of variability in a value or range, for example, within 10%, 5%, or 1% of the limit of the value or range.

[0029] Implementations of this technology relate to a method for the oxidative dehydrogenation (ODH) of ethane to ethylene under conditions that limit acetic acid production. Typically, the oxidative dehydrogenation of ethane involves feeding a gas stream containing ethane and oxygen, and optionally an inert diluent, into an oxidative dehydrogenation reactor including an oxidative dehydrogenation catalyst. Contact between ethane and oxygen and the oxidative dehydrogenation catalyst results in the formation of ethylene and various byproducts, including acetic acid. The generation of byproducts (such as acetic acid) necessitates the costly downstream separation of acetic acid from the target product, ethylene. One object of this disclosure is to limit the extent of acetic acid production during the oxidative dehydrogenation of ethane by including acetic acid along with ethane and oxygen in the initial feed to the method. Reducing the amount of acetic acid produced reduces the complexity and size of the downstream acetic acid separation infrastructure.

[0030] ODH method This disclosure provides a method for the oxidative dehydrogenation of ethane. The method involves contacting a feed stream comprising ethane, oxygen, acetic acid, and optionally an inert diluent with an oxidative dehydrogenation catalyst in an oxidative dehydrogenation reactor under oxidative dehydrogenation conditions to produce a product stream comprising ethylene, unreacted ethane, water, acetic acid, and possibly one or both of carbon dioxide and carbon monoxide. The contribution of acetic acid to the feed stream is from 0.5 vol% to 10 vol%. Figure 1 The system 100 can realize the oxidative dehydrogenation method.

[0031] Feed composition In one aspect, this disclosure seeks to limit acetic acid production in a process for the oxidative dehydrogenation of ethane by including acetic acid in feed stream 101. Acetic acid selectivity using a conventional feed stream without acetic acid can vary considerably depending on the process conditions and the nature of the oxidative dehydrogenation catalyst, but experienced operators should be able to limit the acetic acid selectivity to less than 10% of the converted ethane. By including acetic acid in feed stream 101, the selectivity for acetic acid formation can be reduced. It has been found that by including 0.5 vol% to 10 vol% acetic acid in feed stream 101, acetic acid formation is partially or completely inhibited. It has also been found that negative selectivity for acetic acid is possible, wherein the amount of acetic acid in product stream 103 is reduced compared to the amount of acetic acid in feed stream 101. In some embodiments, the amount of acetic acid in feed stream 101 is 1 vol% to 10 vol%, or 2 vol% to 5 vol%. In some embodiments, the amount of acetic acid in feed stream 101 is about 0.5 vol%, about 1 vol%, about 1.5 vol%, about 2 vol%, about 2.5 vol%, about 3 vol%, about 3.5 vol%, about 4 vol%, about 4.5 vol%, about 5 vol%, about 6 vol%, about 7 vol%, about 8 vol%, about 9 vol%, or about 10 vol%. In some embodiments, the amount of acetic acid in feed stream 101 is about 3 vol%.

[0032] Methods for the oxidative dehydrogenation of ethane fall within the skill of those skilled in the art. The feed preferably comprises an O2:C2H6 ratio outside the flammability limits to prevent process disruption. Users can determine how much inert diluent (e.g., carbon dioxide and water (in vapor form) or a mixture thereof) can be added to ensure the mixture is outside the flammability limits. It should be noted that a feed stream 101 without an inert diluent (though possible) is undesirable because a very small or extremely high O2:C2H6 ratio would be required to maintain it outside the flammability limits. Using vapor as an inert diluent offers the advantage of easier separation from the gaseous target product, but it is also known to increase the selectivity for acetic acid formation. Including acetic acid in the feed stream 101 allows the use of vapor as an inert diluent and easier separation from the target product while minimizing or even avoiding increases in the selectivity for acetic acid formation.

[0033] In some embodiments, the O2:C2H6 volume ratio in feed stream 101 is 0.2:1 to 1:1, 0.3:1 to 0.8:1, or 0.4:1 to 0.7:1. In some embodiments, the contribution of ethane in feed stream 101 is 10 vol% to 80 vol%, 12 vol% to 50 vol%, or 15 vol% to 30 vol%. In some embodiments, the contribution of oxygen in feed stream 101 is 1 vol% to 30 vol%, 5 vol% to 25 vol%, or 8 vol% to 18 vol%.

[0034] Before being introduced into the oxidative dehydrogenation reactor 110, the components of the feed stream 101 can be premixed, or the components can be added separately to the oxidative dehydrogenation reactor 110. It is also possible to premix some components and feed others separately into the oxidative dehydrogenation reactor 110. For example, ethane can be saturated with an inert diluent and introduced into the reactor 110, while oxygen is added separately. The ethane saturated with the inert diluent is then combined with oxygen to form the feed stream 101 that contacts the catalyst. The method also considers adding components to the gas stream in stages, with each stage contributing another component to the gas stream, where the feed stream 101 is formed after the addition of the last component. In the oxidative dehydrogenation method described herein, in addition to ethane, oxygen, and optionally one or more inert diluents, the feed stream 101 introduced into the oxidative dehydrogenation reactor 110 also contains acetic acid.

[0035] Acetic acid may be added alone to the oxidative dehydrogenation reactor 110, or it may be mixed with one or more of ethane, oxygen, or an inert diluent. Acetic acid may be added as glacial acetic acid or in diluted form to provide an amount of acetic acid in the range of about 0.5 to 10.0 vol% in the feed stream 101. The use of dilute aqueous acetic acid is ideal, as water is known to be suitable as an inert diluent in oxidative dehydrogenation processes.

[0036] At a minimum, feed stream 101 can be heated to a temperature above its dew point to ensure all components are in a gaseous state before contacting the oxidative dehydrogenation catalyst. This is particularly relevant when water is used as an inert diluent, as the oxidative dehydrogenation catalyst may be sensitive to liquid water but not to vapor. Components can be heated individually or as a complete mixture. In some embodiments, the temperature of feed stream 101 is at least 150°C, at least 225°C, or at least 300°C.

[0037] Upon entering reactor 110, the temperature of the feed stream 101 or individual components may be below 150°C, provided that the temperature is raised above the dew point before contact with the catalyst. In this case, a portion of reactor 110 may be used to heat the feed stream components to a preferred temperature. This portion of the catalyst bed may be loaded with a thermally conductive non-catalytic material.

[0038] reactor Any known reactor type suitable for hydrocarbon oxidative dehydrogenation can be used with this technology. Conventional fixed-bed reactors are particularly suitable. In a typical fixed-bed reactor, reactants are introduced into the reactor at one end, flow through a stationary catalyst, products are formed, and exit at the other end of the reactor. Designing a suitable fixed-bed reactor follows techniques known for this type of reactor. Those skilled in the art will know what characteristics are required regarding shape and size, reactant input, product output, temperature and pressure control, and the manner in which the catalyst is stationary. In some embodiments, the oxidative dehydrogenation reactor 110 includes a fixed-bed reactor.

[0039] Fluidized bed reactors can also be used. These types of reactors are also well known. Typically, the catalyst is supported by a porous structure or distributor plate located near the bottom of the reactor, and the reactants flow through at a rate sufficient to fluidize the bed (e.g., the catalyst is raised and begins to vortex in a fluidized manner). After contact with the fluidized catalyst, the reactants are converted into products, which are then removed from the top of the reactor. Design considerations include the shape of the reactor and distributor plate, inputs and outputs, and temperature and pressure control, all of which fall within the knowledge of those skilled in the art. In some embodiments, the oxidative dehydrogenation reactor 110 includes a fluidized bed reactor.

[0040] Various tools commonly used in chemical reactors can be used, including flow meters, compressors, valves, and sensors for measuring parameters such as temperature and pressure. Those skilled in the art will likely recognize these components as necessary for operation or compliance with legal obligations related to safety regulations.

[0041] Method conditions The ethane oxidative dehydrogenation method conforming to this disclosure using the oxidative dehydrogenation reactor 110 falls within the scope of knowledge of those skilled in the art. Operators can modify the process conditions and feed composition to optimize product selectivity, conversion, and / or yield.

[0042] As is well known, catalyst beds can have temperature distributions or gradients that vary depending on the reactor type, process conditions, and catalyst composition. Measuring or estimating the temperature of a catalyst bed is also well known in the art, including measuring the temperature at single or multiple points within the catalyst bed. If the temperature variation within the catalyst bed is minimal, ranging from no more than 25°C, preferably no more than 10°C, the temperature can be measured at a single point. Preferably, the catalyst temperature is calculated using the weight-averaged bed temperature from three or more points within the bed. In some embodiments, the oxidative dehydrogenation of ethane can be carried out at temperatures of 300°C to 450°C, 315°C to 425°C, or 330°C to 400°C. In some embodiments, the oxidative dehydrogenation of ethane can be carried out at temperatures of about 300°C, about 315°C, about 330°C, about 400°C, about 425°C, or about 450°C.

[0043] The operating pressure can also be controlled by the operator, including the inlet pressure at which the feed stream is introduced into the oxidative dehydrogenation reactor, where the inlet pressure is higher than the outlet pressure due to the pressure drop along the length of the catalyst bed. The pressures described are relative to the inlet pressure. In some embodiments, the oxidative dehydrogenation of ethane can be carried out at pressures of 0.5 to 100 psig (3.447 to 689.47 kPag) or 15 to 50 psig (103.4 to 344.73 kPag). In some embodiments, the oxidative dehydrogenation of ethane can be carried out at pressures of about 0.5 psig, about 10 psig, about 15 psig, about 20 psig, about 30 psig, about 50 psig, about 75 psig, or about 100 psig.

[0044] In some embodiments, the residence time of ethane in reactor 110 is 0.002 to 30 seconds or 1 to 10 seconds. In some embodiments, the residence time of ethane in reactor 110 is about 0.002 seconds, about 1 second, about 2 seconds, about 3 seconds, about 4 seconds, about 5 seconds, about 6 seconds, about 7 seconds, about 8 seconds, about 9 seconds, about 10 seconds, about 15 seconds, about 20 seconds, about 25 seconds, or about 30 seconds. The flow of reactants and inert diluent can be described in a variety of ways known in the art. Typically, the flow is described and measured using the volume of all feed gases (reactants and diluent) passing through the active catalyst bed volume over one hour, or the gas hourly space velocity (GHSV). In some embodiments, the GHSV is 500 to 30,000 h⁻¹. -1 1000h -1 Up to 20000h -1 1500h -1 Up to 10000h -1 or 2000h -1 Up to 10000h -1 In some implementations, the GHSV is approximately 500h. -1 Approximately 1000 hours -1 Approximately 1500 hours -1 Approximately 2000 hours -1 Approximately 3000 hours -1 Approximately 4000 hours -1 Approximately 5000 hours -1 Approximately 6000 hours -1 Approximately 7000 hours -1 Approximately 8000 hours -1 Approximately 9000 hours -1 Approximately 10,000 hours -1 Approximately 15,000 hours -1 Approximately 20,000 hours -1 Approximately 25,000 hours -1 or approximately 30,000 hours -1 .

[0045] Flow rate can also be measured as weight time space velocity (WHSV), which, relative to volume, describes the flow as the weight of gas flowing through the active catalyst bed per hour. When calculating WHSV, the weight of the gas may include only the reactants, but may also include the diluent added to the gas mixture. In some embodiments, the WHSV (including the weight of the diluent) is 0.5 h⁻¹. -1 up to 50h -1 1.0 to 25.0h -1 Or 2.0 to 10.0h -1 In some implementations, the WHSV (including the weight of the diluent) is approximately 0.5 h. -1 Approximately 1.0h -1 Approximately 1.5 hours -1 Approximately 2.0 hours -1 Approximately 3.0 hours -1 Approximately 4.0 hours -1 Approximately 5.0 hours -1 Approximately 6.0 hours -1 Approximately 7.0 hours -1 Approximately 8.0 hours -1 Approximately 9.0 hours -1 Approximately 10 hours -1 Approximately 15 hours -1 Approximately 20 hours -1 Approximately 25 hours -1 Approximately 30 hours -1 Approximately 40 hours -1 or about 50 hours -1 .

[0046] The flow of feed stream 101 through reactor 110 can also be described as the linear velocity (cm / s) of the feed stream, defined in the art as the feed stream velocity 101 / cross-sectional surface area of ​​reactor 110 / void fraction of catalyst bed. The velocity generally refers to the total velocity of all gases entering reactor 110 and is measured at the point when oxygen and alkanes first contact the catalyst, and at the temperature and pressure at that point. The cross-section of reactor 110 is also measured at the inlet of the catalyst bed. The void fraction of the catalyst bed is defined as the volume of voids in the catalyst bed / the total volume of the catalyst bed. The volume of voids refers to the voids between catalyst particles, excluding the pore volume within the catalyst particles. In some embodiments, the linear velocity is from 5 cm / sec to 1500 cm / sec, 10 cm / sec to 500 cm / sec, or 25 cm / sec to 350 cm / sec. In some embodiments, the linear velocity is approximately 5 cm / sec, approximately 10 cm / sec, approximately 15 cm / sec, approximately 20 cm / sec, approximately 25 cm / sec, approximately 50 cm / sec, approximately 75 cm / sec, approximately 100 cm / sec, approximately 150 cm / sec, approximately 200 cm / sec, approximately 250 cm / sec, approximately 300 cm / sec, approximately 350 cm / sec, approximately 400 cm / sec, approximately 500 cm / sec, approximately 600 cm / sec, approximately 700 cm / sec, approximately 800 cm / sec, approximately 900 cm / sec, approximately 1000 cm / sec, approximately 1250 cm / sec, or approximately 1500 cm / sec. In some embodiments, the space-time yield (productivity) of ethylene, expressed in g ethylene / hour per kg catalyst, is at least 900 h⁻¹. -1 Or at least 1500h -1 In some implementations, the space-time yield (productivity) of ethylene, expressed in g ethylene / hour per kg catalyst, is at least 3500 h at temperatures between 350°C and 400°C. -1 It should be noted that as the temperature increases, the productivity of the catalyst increases, but this is usually accompanied by a decrease in the selectivity for ethylene production.

[0047] Optimization or adjustment of method conditions can affect the conversion and corresponding selectivity of ethane, including selectivity to ethylene and acetic acid. In some embodiments, the method has a selectivity to ethylene of at least 60%, at least 75%, or at least 90%. In some embodiments, the conversion of ethane is at least 25%, at least 40%, at least 50%, at least 60%, or at least 70%.

[0048] catalyst Many known catalysts are available for the oxidative dehydrogenation of ethane. Mixed metal oxides containing molybdenum and vanadium are particularly suitable for this purpose. Typically, oxidative dehydrogenation catalysts comprise mixed metal oxide catalysts selected from the following: i) Catalysts of the following formula: Mo a V b Te c Nb d Pd e O f Where: a, b, c, d, e, and f are the relative atomic weights of elements Mo, V, Te, Nb, Pd, and O, respectively; and when a = 1, b = 0.01 to 1.0, preferably 0.1 to 0.4, c = 0.01 to 1.0, preferably 0.1 to 0.3, d = 0.01 to 1.0, preferably 0.1 to 0.3, and 0.00 ≤ e ≤ 0.10, preferably 0.03 to 0.1, f is the number of catalyst valence states. ii) Catalysts of the following formula: Mo a E k G l O f Wherein: E is selected from Ba, Ca, Cr, Mn, Nb, Ta, Ti, Te, V, W and mixtures thereof; G is selected from Bi, Ce, Co, Cu, Fe, K, Mg, V, Ni, P, Pb, Sb, Si, Sn, Ti, U and mixtures thereof; a=1; k is 0 to 2, preferably 0.2 to 0.6; l=0 to 2, preferably 0.2 to 0.6, provided that for C O The total value of l for Ni, Fe and their mixtures is less than 0.5; and f is the number of catalyst valence states. iii) Catalysts of the following formula: V m Mo n Nb o Te p Me q O f Where: Me is a metal selected from Ta, Ti, W, Hf, Zr, Sb and mixtures thereof; m is 0.1 to 3, and in some cases 0.5 to 1; n is 0.5 to 1.5, and in some cases 0.5 to 1; o is 0.001 to 3, and in some cases 0.01 to 1; p is 0.001 to 5, and in some cases 0.01 to 1; q is 0 to 2, and in some cases 0.01 to 1; and f is the number of catalyst valence states; and iv) Catalysts of the following formula: Mo a Vr X st Z u M v O f Wherein: X is at least one of Nb and Ta; Z is at least one of Te, Ga, Pd, W, Bi, and Al, and in some embodiments, at least one of Te, Pd, W, and Bi; M is at least one of Fe, Co, Cu, Cr, Ti, Ce, Zr, Mn, Pb, Mg, Sn, Pt, Si, La, K, Ag, and In, and in some cases, at least one of Fe, Co, Cu, Cr, Ti, Ce, Zr, Mn, Mg, Sn, Pt, La, Ag, and In. One less; a = 1.0 (normalized); r = 0.05 to 1.0, in some embodiments, 0.05 to 0.5; s = 0.001 to 1.0, in some embodiments, 0.01 to 0.4; t = 0.001 to 1.0, in some embodiments, 0.01 to 0.4; u = 0.001 to 0.5, in some embodiments, 0.01 to 0.03; v = 0.001 to 0.3, in some embodiments, 0.01 to 0.2; and f is the number of catalyst valence states.

[0049] In some embodiments, the catalyst is a catalyst comprising molybdenum, vanadium, tellurium, niobium and oxygen, wherein the molar ratio of molybdenum to vanadium is 1:0.12 to 1:0.49, the molar ratio of molybdenum to tellurium is 1:0.01 to 1:0.30, the molar ratio of molybdenum to niobium is 1:0.01 to 1:0.30, and oxygen is present in an amount satisfying the valence of any present metal oxide.

[0050] In some embodiments, the catalyst is a catalyst comprising molybdenum, vanadium, tellurium, niobium and oxygen, wherein the molar ratio of molybdenum to vanadium is 1:0.20 to 1:0.45, the molar ratio of molybdenum to tellurium is 1:0.05 to 1:0.25, the molar ratio of molybdenum to niobium is 1:0.05 to 1:0.25, and oxygen is present in an amount satisfying the valence of any present metal oxide.

[0051] In some embodiments, the catalyst is a catalyst comprising molybdenum, vanadium, tellurium, niobium and oxygen, wherein the molar ratio of molybdenum to vanadium is 1:0.25 to 1:0.40, the molar ratio of molybdenum to tellurium is 1:0.10 to 1:0.20, the molar ratio of molybdenum to niobium is 1:0.10 to 1:0.20, and oxygen is present in an amount satisfying the valence of any present metal oxide.

[0052] In some embodiments, the catalyst is a catalyst comprising molybdenum, vanadium, tellurium, niobium and oxygen, wherein the molar ratio of molybdenum to vanadium is 1:0.30 to 1:0.35, the molar ratio of molybdenum to tellurium is 1:0.13 to 1:0.17, the molar ratio of molybdenum to niobium is 1:0.12 to 1:0.14, and oxygen is present in an amount satisfying the valence of any present metal oxide.

[0053] Recycled acetic acid In another aspect of this disclosure, acetic acid produced in the ethane oxidative dehydrogenation process can be recovered and recycled back into the feed stream. The product stream can be subjected to processing steps to separate the target product, ethylene. As described in the art, the first processing step ideally involves removing acetic acid and water from the product stream, typically including cooling the product stream to condense a substantial portion of the acetic acid and vapor, which can then be readily separated from the gaseous components as a liquid stream 107. The gaseous components forming a portion of the gaseous component stream 105 contain ethane, ethylene, and carbon dioxide, which can then be subjected to further processing steps, which may include separating carbon dioxide from ethane and ethylene, and subsequently separating ethane from ethylene. The separated ethane can be recycled to form a portion of the feed stream 101. Cooling and separation of acetic acid from the product stream 103 can be non-dilutive, for example, by passing the product stream 103 through a heat exchanger. Cooling and separation of acetic acid from the product stream 103 can be dilutive, for example, by introducing the product stream 103 into a quench tower 118, where cold water is added to stream 103. Combinations of methods may also be employed. Regardless of the method used for separation (dilutive, non-dilutive, or a combination of both), the liquid stream 107, including aqueous acetic acid, can be recycled for feed stream 101.

[0054] The concentration of acetic acid in liquid stream 107 can vary depending on the initial concentration in product stream 103 and the method or combination of methods used for separation. For example, dilutive cooling using quench tower 118 will result in a much lower concentration of acetic acid due to the addition of water during quenching. Those skilled in the art will be able to determine the concentration of acetic acid in liquid stream 107 and then infer how much of liquid stream 107 can be recycled to feed stream 101 to provide 0.5 to 10 vol% acetic acid. Where it is not necessary for all of liquid stream 107 to provide the desired vol% acetic acid in feed stream 101, a fraction of liquid stream 107 can be recycled for addition to feed stream 101, and the remainder can be sent for further processing, for example, upgrading to glacial acetic acid using methods known in the art. Where the concentration of acetic acid in liquid stream 107 is insufficient to provide enough acetic acid to provide 0.5 to 10 vol% in the feed stream, the difference can be made up using an additional acetic acid source. The gaseous portions of streams 101, 103, and 105 have compositions outside their respective flammability limits.

[0055] In some implementations, multiple implementations of the ODH reactor 110 are used to perform oxidative dehydrogenation of hydrocarbons (examples shown in...). Figure 2A and 2B (Middle). The reactor can be in a parallel flow configuration, a series flow configuration, or a combination of these. For example... Figure 2A As shown, oxidative dehydrogenation of hydrocarbons can be performed via system 200a, which includes two oxidative dehydrogenation reactors 210A and 210B in a series flow configuration. Each of ODH reactors 210A and 210B is an implementation of the aforementioned ODH reactor 110. Feed stream 101 flows to ODH reactor 210A and contacts the catalyst in ODH reactor 210A.

[0056] Product stream 203a flows from ODH reactor 210A to quench tower 218A, where acetic acid is separated from product stream 203a. Liquid stream 207a, containing acetic acid, flows out from quench tower 218A. Gaseous component stream 205a flows from quench tower 218A to ODH reactor 210B and contacts the catalyst in ODH reactor 210B. In some embodiments, makeup stream 251 flows to ODH reactor 210B and contacts the catalyst in ODH reactor 210B. In some embodiments, makeup stream 251 has the same composition as feed stream 101. In some embodiments, makeup stream 251 has the same composition as feed stream 101, but the amount of each component is different from that of feed stream 101. In some embodiments, the composition and flow rate of makeup stream 251 are determined such that the entire stream entering ODH reactor 210B (gaseous component stream 205a and makeup stream 251 together) has the same composition as feed stream 101. Before introducing components into ODH reactor 210B, the components of gaseous component stream 205a and makeup stream 251 can be premixed, or the components can be added separately to ODH reactor 210B. Alternatively, some components can be premixed, while others are fed separately into ODH reactor 210B.

[0057] Product stream 203b flows from ODH reactor 210B to quench tower 218B, where acetic acid is separated from product stream 203b. Liquid stream 207b, containing acetic acid, flows out of quench tower 218B. Gaseous component stream 205b flows out of quench tower 210B. The composition of feed stream 101 and makeup stream 251 is adjustable such that the total stream entering each ODH reactor 210A, 210B has an acetic acid content ranging from 0.5 vol% to 10 vol%. The composition of feed stream 101 and makeup stream 251 is also adjustable such that the gaseous portions of streams 101, 203a, 203b, 205a, 205b, and 251 have a composition outside their respective flammability limits. In some embodiments, the composition of the feed stream 101, the operating temperature of the ODH reactor 210A, the operating pressure of the ODH reactor 210A, the composition of the makeup stream 251, the amount of the gaseous component stream 205a diverted and bypassing the ODH reactor 210B, or a combination thereof, are adjusted such that the ethylene content of the total stream entering the ODH reactor 210B is equal to or less than 20% by weight.

[0058] In some embodiments, a portion of the gaseous component stream 205a is diverted and bypasses the ODH reactor 210B and the quench tower 218B, and is combined with the gaseous component stream 205b. The gaseous component stream 205b (and in some cases, a portion of the gaseous component stream 205a already combined with 205b), containing ethane, ethylene, and carbon dioxide, may then undergo further processing steps, which may include separating carbon dioxide from ethane and ethylene, and subsequently separating ethane from ethylene. In some embodiments, liquid streams 207a, 207b, or both 207a and 207b are recirculated to form at least a portion of the feed stream 101. In some embodiments, liquid streams 207a, 207b, or both 207a and 207b are recirculated to form at least a portion of the makeup stream 251.

[0059] The operating conditions (temperature and pressure) of ODH reactors 210A and 210B can be adjusted to improve product distribution control, increase net ethylene yield, reduce the amount of dilution gas used in the ODH process, or any combination thereof. For example, upstream ODH reactor 210A can operate at a lower temperature compared to downstream ODH reactor 210B, a higher temperature compared to downstream ODH reactor 210B, or the same temperature as downstream ODH reactor 210B. Similarly, upstream ODH reactor 210A can operate at a lower pressure compared to downstream ODH reactor 210B, a higher pressure compared to downstream ODH reactor 210B, or the same pressure as downstream ODH reactor 210B. In an embodiment where upstream ODH reactor 210A operates at a lower pressure (or the same pressure) compared to downstream ODH reactor 210B, system 200b may include an ejector 250, such as... Figure 2B As shown in the diagram. In such an embodiment, the makeup stream 251 can be used as the kinetic fluid via the ejector 250. The makeup stream 251 and the gaseous component stream 205a are mixed as they flow through the ejector 250, and the mixture flows from the ejector 250 to the ODH reactor 210B. The operating conditions of the kinetic fluid (makeup stream 251) to the ejector 250 can be adjusted accordingly to meet the target operating conditions of the ODH reactor 210B. Similarly, the operating conditions of the feed stream 101 can be adjusted accordingly to meet the target operating conditions of the ODH reactor 210A.

[0060] refer to Figure 3A Method 300a can be implemented for an oxidative dehydrogenation process. For example, method 300a can be implemented by any system 100, 200a, or 200b. In step 302, in an oxidative dehydrogenation reactor (e.g., ODH reactor 110, 210A, or 210B), a feed stream (e.g., feed stream 101) is contacted with an oxidative dehydrogenation catalyst under oxidative dehydrogenation conditions. As mentioned above, feed stream 101 contains ethane, oxygen, and acetic acid. The concentration of acetic acid in feed stream 101 is 0.5 to 10% by volume. In some cases, feed stream 101 contains additional components, such as an inert diluent. Contacting feed stream 101 with the oxidative dehydrogenation catalyst in an ODH reactor (110, 210A, 210B) under oxidative dehydrogenation conditions produces product streams (e.g., product streams 103, 203a, or 203b). As mentioned above, the product streams (103, 203a, 203b) contain ethylene, unreacted ethane, water, and acetic acid. In some cases, the product streams (103, 203a, 203b) may contain additional components, such as carbon monoxide, carbon dioxide, or both.

[0061] refer to Figure 3BMethod 300b can be implemented for an oxidative dehydrogenation process. For example, method 300b can be implemented by any system 200a or 200b. Similar to method 300a, method 300b includes step 302. In step 304, product stream 203a is separated into a liquid stream (e.g., liquid stream 207a) and a gaseous component stream (e.g., gaseous component stream 205a). Step 304 can be implemented by a downstream separation method, such as a quench tower 218A. In some embodiments, at least a portion of liquid stream 207a is recycled as part of feed stream 101 to ODH reactor 210A.

[0062] In step 306, in the second oxidative dehydrogenation reactor (e.g., ODH reactor 210B), the makeup stream (e.g., makeup stream 251) and the gaseous component stream 205a are contacted with the second oxidative dehydrogenation catalyst under oxidative dehydrogenation conditions. The makeup stream 251 and the gaseous component stream 205a are contacted with the oxidative dehydrogenation catalyst in the ODH reactor 210B under oxidative dehydrogenation conditions to produce a second product stream (e.g., product stream 203b). As mentioned above, product stream 203b contains ethylene, unreacted ethane, water, and acetic acid. In some cases, product stream 203b contains additional components, such as carbon monoxide, carbon dioxide, or both. The total concentration of acetic acid in the makeup stream 251 and the gaseous component stream 205a (together) entering the ODH reactor 210B is 0.5 to 10% by volume. In step 308, product stream 203b is separated into a second liquid stream (e.g., liquid stream 207b) and a second gaseous component stream (e.g., gaseous component stream 205b). Step 308 can be achieved via a second downstream separation method, such as quench tower 218B. In some embodiments, at least a portion of the second liquid stream 207b is recycled to ODH reactor 210A as part of feed stream 101, or to second ODH reactor 210B as part of makeup stream 251. Example

[0063] The following examples are merely illustrative of the subject matter of this disclosure and are not intended to be limiting. Computational modeling using the ODH method demonstrates the effect of adding acetic acid to the feed stream on acetic acid production. The modeling is based on experimental data generated using catalysts containing molybdenum, vanadium, niobium, and tellurium. The model demonstrates the effect of adding acetic acid to the feed stream under ODH method conditions of different temperatures, GHSV, and feed compositions. For each method condition, a baseline case of 0 vol% acetic acid in the feed stream is compared with cases of 2 vol%, 5 vol%, and 10 vol% acetic acid in the feed.

[0064] Cases 1-5 are modeling examples, where each case represents different method conditions or feed compositions, varying and modeled for 0 vol%, 2 vol%, 5 vol%, and 10 vol% acetic acid in the feed stream. Table 1 summarizes the method conditions and feed compositions for each of cases 1-5. The feed stream compositions only list O2 vol%, C2H6 vol%, and CO2 vol%, with the remainder including the addition of water and acetic acid to reach a total of 100 vol%. Table 2-1 summarizes the results, including the selectivity of ethylene and acetic acid for 0 vol%, 2 vol%, 5 vol%, and 10 vol% acetic acid added to the feed stream, as well as the ethane conversion (%), showing a decrease in acetic acid selectivity for each vol% acetic acid in the feed stream compared to when there is no acetic acid in the feed stream. The results for 0.5 vol% acetic acid in the feed stream, including ethane conversion and the selectivity of ethylene, acetic acid, and carbon monoxide / carbon dioxide, are summarized in Table 2-2 as the intermediate values ​​between 0 vol% and 2 vol% in the feed stream shown in Table 2-1. Table 2-2 confirms that as the acetic acid content in the feed stream increases, the selectivity of acetic acid in the product tends to decrease.

[0065] Table 1: Method conditions and feed composition for each of scenarios 1 to 5. (1) Note that commercially available ethane often contains impurities or other light hydrocarbons, such as methane and propane, up to 5% by volume. These components are considered part of the ethane composition and are not included in the volume of each component in the determined feed stream.

[0066] Table 2-1: Catalyst performance for each of cases 1 to 5. Table 2-2: Method conditions and catalyst performance for each of cases 1 to 5. Experiments were conducted using a fixed-bed reactor unit (FBRU) apparatus to add acetic acid to the feed stream for the oxidative dehydrogenation of ethane. The FBRU apparatus comprises two vertically oriented fixed-bed tubular reactors connected in series. Each reactor is an SS316L tube with an outer diameter of 1'' and a length of 34" and is encased in an electrically heated jacket and sealed with ceramic insulation. Each reactor contains an identical catalyst bed consisting of 1 unit of catalyst / 2.14 units of DENSTONE® 99 (primarily α-alumina) powder. The total weight of the catalyst in each reactor is 143 g, and the catalyst has the formula MoV0. 0.40 Nb 0.16 Te 0.14 O, relative to Mo, is represented by a 1:1 ratio. The relative atomic weights of each component are shown by subscripts. Below and above the catalyst bed, the remainder of the reactor is filled with quartz powder, which is secured in place with glass wool to minimize the risk of bed movement during experimental runs.

[0067] The temperature of each reactor is monitored using seven thermocouples located at corresponding points within each reactor, four of which are situated within the respective catalyst beds. Temperature control (especially at lower temperatures) is limited and results in fluctuations. The temperatures listed in the example represent the average temperature at eight different locations within the two catalyst beds. The temperatures of both reactors are controlled by adjusting the pressure within the water jacket surrounding each reactor and the boiling point of the water.

[0068] For the experimental setup, the pressure at the inlet was monitored using a pressure sensor located directly upstream of the first reactor. The product stream exiting the second reactor was routed through a condenser unit and then discharged into the air, indicating that the pressure at this point was close to 0 psig.

[0069] Cases 6 and 7 are experimental examples conducted using an FBRU device, demonstrating a practical reduction in the selectivity for acetic acid production within the physical settings. Table 3 summarizes the method conditions and feed composition for each of cases 6 and 7. The feed composition lists only O2, C2H6, CO2, and acetic acid, with the remainder including water required to reach a total volume percentage of 100%. Table 4 summarizes the results, including the selectivity for ethylene and acetic acid, as well as the conversion rate (%) of ethane, showing that, similar to modeling cases 1-5, the addition of acetic acid to the feed stream leads to a decrease in acetic acid selectivity.

[0070] Table 3: Method conditions and feed composition for scenarios 6 and 7. (1) Note that commercially available ethane often contains impurities or other light hydrocarbons, such as methane and propane, up to 5% by volume. These components are considered part of the ethane composition and are not included in the volume of each component in the determined feed stream.

[0071] Table 4: Catalyst performance in scenarios 6 and 7. For scenarios 8-11, a similar FBRU apparatus was used. For scenarios 8-11, reactor temperature, inlet pressure, GHSV, ethane feed volume fraction, and oxygen feed volume fraction were kept constant. Each scenario in 8-11 included two experiments, where the first experiment ('a') contained both vapor and acetic acid in the feed, and the second experiment ('b') contained vapor in the feed but no acetic acid. The reactor operating conditions and feed composition (volume %) for scenarios 8-11 are provided in Table 5. The catalyst activity for scenarios 8-11 is provided in Table 6. Table 7 shows the changes in catalyst activity and product distribution with acetic acid in the feed ('a') compared to scenario 8-11 with no acetic acid in the feed ('b').

[0072] The following summarizes the trends observed in scenarios 8-11. The presence or absence of acetic acid in the feed appears to have no effect on ethane conversion. The selectivity for ethylene remains the same or increases when acetic acid is present in the feed ('a') compared to when acetic acid is absent in the feed ('b'). The selectivity for carbon monoxide / carbon dioxide increases when acetic acid is present in the feed ('a') compared to when acetic acid is absent in the feed ('b'). The selectivity for acetic acid decreases or is completely inhibited when acetic acid is present in the feed ('a') compared to when acetic acid is absent in the feed ('b').

[0073] Compared to scenario 8, scenario 9 was conducted at an increased temperature of 359°C. A comparison of scenarios 8 and 9 shows that with increasing temperature, the change in acetic acid selectivity is negligible (Table 7), while ethylene selectivity decreases and carbon monoxide / carbon dioxide selectivity increases (Table 7), and ethane conversion increases with increasing temperature (Table 6). Compared to scenario 8, scenario 10 was conducted at increased pressure. A comparison of scenarios 8 and 10 shows that with increasing pressure, the decrease in acetic acid selectivity decreases (Table 7), the increase in ethylene selectivity is almost completely suppressed with increasing pressure (Table 7), the increase in carbon monoxide / carbon dioxide selectivity increases with increasing pressure (Table 7), and ethane conversion increases with increasing pressure (Table 6).

[0074] These experiments suggest that a portion of acetic acid (when present in the feed) is being converted to ethylene, carbon monoxide, and / or carbon dioxide, and that the selectivity can be altered by reducing the temperature, pressure, or both, to favor ethylene over carbon monoxide / carbon dioxide. However, reducing the temperature, pressure, or both may also decrease ethane conversion. In some cases, reducing the operating pressure can have a greater impact on altering the selectivity towards ethylene over carbon monoxide / carbon dioxide than reducing the operating temperature. Using multiple reactors (e.g., in systems 200a and 200b) allows the ODH process to occur at various combinations of operating temperatures and pressures to meet both the target ethane conversion and the target ethylene selectivity.

[0075] Table 5: Method conditions and feed composition for each of scenarios 8 to 11 Table 6: Feed composition and catalyst performance for each of scenarios 8 to 11. Table 7: Comparison of catalyst performance for each of scenarios 8 to 11. In scenario 12, the amounts of ethylene and ethane differ in the three experiments. The individual amounts of ethylene and ethane differ, while the total amount of ethylene and ethane in the feed remains constant. The amounts of other components (water, oxygen, and carbon monoxide) remain constant. For all experiments in scenario 12, the reactor inlet pressure is ambient pressure, and the WHSV is 1.29–1.31 h⁻¹. -1 GHSV is 1318-1322 h -1 The reactor temperature was 339-342℃. Table 8 provides the feed composition, ethane conversion, and product selectivity. Results from scenario 12 show that increasing the ethylene composition from 11 wt% to 28 wt% resulted in increased selectivity towards carbon monoxide / carbon dioxide and acetic acid, decreased selectivity towards ethylene, and no observable trend in ethane conversion.

[0076] Table 8: Feed composition and catalyst performance in scenario 12. Industrial applicability This disclosure relates to a method for the oxidative dehydrogenation of ethane. The method is suitable for limiting the amount of acetic acid produced by including acetic acid along with ethane and oxygen as part of the feed to an ODH process.

Claims

1. A method for the oxidative dehydrogenation of ethane, the method comprising: A feed stream containing ethane, oxygen, and acetic acid is contacted with an oxidative dehydrogenation catalyst in an oxidative dehydrogenation reactor under oxidative dehydrogenation conditions to produce a product stream containing ethylene, unreacted ethane, water, and acetic acid, wherein the concentration of acetic acid in the feed stream is 0.5 to 10% of the feed stream volume.

2. The method of claim 1, wherein the concentration of acetic acid in the feed stream is 2 to 5% of the feed stream volume.

3. The method according to claim 1, wherein the concentration of acetic acid in the feed stream is greater than 2% to 10% of the feed stream volume.

4. The method according to any one of claims 1 to 3, the method further comprising a downstream separation method, wherein the product stream is separated into a liquid stream comprising water and acetic acid, and a gaseous component stream comprising ethylene and unreacted ethane, and at least a portion of the liquid stream is recycled back to the reactor as part of the feed stream.

5. The method of claim 4, wherein the liquid stream is diluted with water to achieve a desired amount of acetic acid in the feed stream.

6. The method of claim 4, wherein the fractionation rate of the liquid flow is adjusted to achieve a desired amount of acetic acid in the feed stream.

7. The method according to any one of claims 1 to 3, wherein the feed stream comprises: i) 0.5 to 10% by volume acetic acid; ii) Oxygen and ethane in a molar ratio of 0.5 to 0.7; and iii) The molar ratio of H2O and CO2 in the feed composition is outside the flammability limit.

8. The method according to any one of claims 1 to 3, wherein the reactor operates at a temperature of 300°C to 425°C.

9. The method according to any one of claims 1 to 3, wherein the reactor operates at a temperature of 315°C to 400°C.

10. The method according to any one of claims 1 to 3, wherein the reactor operates at a pressure of 0.5 psig to 100 psig.

11. The method according to any one of claims 1 to 3, wherein the reactor operates at a pressure of 15 psig to 50 psig.

12. The method according to any one of claims 1 to 3, wherein the gas hourly space velocity of the product stream is 500 h⁻¹. -1 Up to 30000h -1 .

13. The method according to any one of claims 1 to 3, wherein the gas hourly space velocity of the product stream is 1000 h⁻¹. -1 Up to 15000h -1 .

14. The method according to any one of claims 1 to 3, wherein the gas hourly space velocity of the product stream is 500 h⁻¹. -1 Up to 4000h -1 .

15. The method according to any one of claims 1 to 3, wherein the catalyst comprises one or more catalysts selected from the group consisting of: i) Catalysts of the following formula: For a V b The c Nb d Pd e Oh f in: a, b, c, d, e, and f are the relative atomic weights of elements Mo, V, Te, Nb, Pd, and O, respectively; and when a = 1, b = 0.01 to 1.0, c = 0.01 to 1.0, d = 0.01 to 1.0, and 0.00 ≤ e ≤ 0.10, f is the number of valence states that satisfy the catalyst. ii) Catalysts of the following formula: Mon a E k G l O f Where: E is selected from Ba, Ca, Cr, Mn, Nb, Ta, Ti, Te, V, W and their mixtures; G is selected from Bi, Ce, Co, Cu, Fe, K, Mg, V, Ni, P, Pb, Sb, Si, Sn, Ti, U and their mixtures; a=1; k is 0 to 2; l=0 to 2, provided that the total value of l for Co, Ni, Fe and their mixtures is less than 0.5; and f is the number of catalyst valence states. iii) Catalysts of the following formula: V m For n Nb o The p I q Oh f Where: Me is a metal selected from Ta, Ti, W, Hf, Zr, Sb and mixtures thereof; m is 0.1 to 3; n is 0.5 to 1.5; o is 0.001 to 3; p is 0.001 to 5; q is 0 to 2; and f is the number of catalyst valence states; and iv) Catalysts of the following formula: Mon a V r X s Y t Z u M v O f Wherein: X is at least one of Nb and Ta; Y is at least one of Sb and Ni; Z is at least one of Te, Ga, Pd, W, Bi and Al; M is at least one of Fe, Co, Cu, Cr, Ti, Ce, Zr, Mn, Pb, Mg, Sn, Pt, Si, La, K, Ag and In; after normalization a = 1.0; r = 0.05 to 1.0; s = 0.001 to 1.0; t = 0.001 to 1.0; u = 0.001 to 0.5; v = 0.001 to 0.3; and f is the number of catalyst valence states.

16. The method of claim 4, wherein the liquid stream comprises less than 10% by volume of acetic acid.

17. The method according to any one of claims 1 to 3, wherein the selectivity for generating ethylene is 75% to 99%.

18. The method according to any one of claims 1 to 3, wherein the selectivity of CO2 generation is equal to or less than 10%.

19. The method according to any one of claims 1 to 3, wherein the selectivity of CO generation is equal to or less than 11%.

20. The method according to any one of claims 1 to 3, wherein the feed stream comprises 2 to 3 vol% acetic acid, 29 to 57 vol% H2O, 16 to 26 vol% C2H6, 8 to 14 vol% O2 and 17 to 28 vol% CO2.

21. The method of claim 4, wherein the makeup stream and the gaseous component stream are contacted with the second oxidative dehydrogenation catalyst in the second oxidative dehydrogenation reactor under oxidative dehydrogenation conditions to produce a second product stream comprising ethylene, unreacted ethane, water and acetic acid, wherein the total concentration of acetic acid in the makeup stream and the gaseous component stream entering the second oxidative dehydrogenation reactor is 0.5 to 10 volumes.

22. The method of claim 21, further comprising a second downstream separation method, wherein the second product stream is separated into a second liquid stream comprising water and acetic acid, and a second gaseous component stream comprising ethylene and unreacted ethane, and at least a portion of the second liquid stream is recycled as part of a feed stream to the second oxidative dehydrogenation reactor, or as part of a makeup stream to the second oxidative dehydrogenation reactor.

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