HK1159075A - Process for catalytically producing ethylene directly from acetic acid in a single reaction zone - Google Patents

Description

Process for the direct catalytic production of ethylene from acetic acid in a single reaction zone

Priority

This application is based on U.S. patent application serial No.12/221,137 filed on 31/7/2008, having the same title, and therefore claims priority thereto and the disclosure of which is incorporated herein by reference.

Technical Field

The present invention generally relates to a process for producing ethylene from acetic acid. More particularly, the present invention relates to processes wherein acetic acid is directly converted to ethylene in a single reaction zone that may include a combination of multiple catalysts. Without wishing to be bound by any theory, it is believed that the catalyst is capable of simultaneously hydrogenating acetic acid and converting intermediates to ethylene with high selectivity and yield.

Background

There is a long felt need for an economically viable process for converting acetic acid to ethylene. Ethylene is an important commercial feedstock for a variety of industrial products; for example, ethylene can then be converted into various polymers and other monomer products. Fluctuating natural gas and crude oil prices contribute to fluctuating costs of conventionally produced petroleum or natural gas derived ethylene, thereby creating a greater demand for alternative sources of ethylene than ever before as oil prices rise.

Ethylene from various ethyl esters was reported to be produced in the gas phase over a zeolite catalyst at temperatures in the range of 150-300 ℃. Types of ethyl esters that can be used include ethyl esters of formic acid, acetic acid, and propionic acid. See, for example, U.S. Pat. No.4,620,050 to Cognion et al, where selectivity is reported to be acceptable.

U.S. Pat. No.4,270,015 to Knifton describes a process involving a two-step process to obtain ethylene: wherein a mixture of carbon monoxide and hydrogen, commonly referred to as synthesis gas, is reacted with a carboxylic acid containing 2 to 4 carbon atoms to form the corresponding ethyl ester of said carboxylic acid, which is subsequently pyrolyzed in a quartz reactor at elevated temperatures of about 200 ℃ to 600 ℃ to obtain ethylene. The ethylene so produced contains other hydrocarbons, particularly ethane as an impurity. It has also been reported that pyrolysis of pure ethyl propionate at 460 ℃ can bring the concentration of ethane to high values, close to 5%. More importantly, very low ester conversion and ethylene yields are reported.

U.S. Pat. No.4,399,305 to Schreck describes the use of DuPont de Nemours&Co under the trade nameThe cracking catalyst, which is composed of a commercially available perfluorosulfonic acid resin, obtains high purity ethylene from ethyl acetate.

On the other hand, ball. Soc. Chim. Belg. by Malinowski et al (1985), 94(2), 93-5 disclose the use of a substrate, such as acetic acid, on a support material such as Silica (SiO) (SiO. RTM. I. B₂) Or titanium dioxide (TiO)₂) The reaction on the lower titanium, which is heterogenized above, produces a product mixture comprising diethyl ether, ethylene and methane, with poor selectivity.

WO 2003/040037 discloses the use of crystalline microporous metalloaluminophosphates (ELAPO), particularly SAPO-type zeolites such as SAPO-5, SAPO-11, SAPO-20, SAPO-18 and SAPO-34, having Si/Al ratios of 0.03 to 017, as adsorbents or catalysts for the production of olefins from oxygenated feedstocks containing methanol, ethanol, n-propanol, isopropanol, C4-C20 alcohols, methyl ethyl ether, dimethyl ether, diethyl ether, diisopropyl ether, formaldehyde, dimethyl carbonate, dimethyl ketone and/or acetic acid. Similar disclosures utilize silicoaluminophosphate molecular sieves comprising at least one intergrown phase of molecular sieve. It is reported that in this process an oxygenate-containing feedstock contacts a catalyst comprising a molecular sieve in a reaction zone of a reactor under conditions effective to produce light olefins, particularly ethylene and propylene. See U.S. patent No.6,812,372 to Janssen. Reference to such an oxygenate feedstock includes acetic acid, but the disclosure appears to be limited to methanol or dimethyl ether. See, also, U.S. Pat. No.6,509,290 to Vaughn, which also discloses converting an oxygenate feedstock to olefins.

Bimetallic ruthenium-tin/silica catalysts by reacting tetrabutyltin with a supportThe preparation is carried out by reaction of ruthenium dioxide on silica. These catalysts are reported to exhibit different selectivities depending on their value (content) of the tin/ruthenium ratio (Sn/Ru). Specifically, the selectivity of the hydrogenolysis of ethyl acetate was reported to be very different depending on the Sn/Ru ratio in the catalyst. For example, for ruthenium alone in SiO₂The reaction is not selective: producing methane, ethane, carbon monoxide, carbon dioxide, and ethanol and acetic acid. However, for low tin contents, the catalyst is reported to have considerable selectivity for the formation of acetic acid, and at the same time at higher Sn/Ru ratios, ethanol is the only product detected. See students in Surface Science and Catalysis (1989), Volume Date 1988, 48(struct. React. surf.), 591-600, Loessard et al.

Catalytic reduction of acetic acid was also investigated. For example, j.chem.res, Synopses (1980), (11), 373, by Hindermann et al, discloses the catalytic reduction of acetic acid on iron and on base-promoted iron. In their studies they found that the reduction of acetic acid on base-promoted iron follows at least two different paths depending on the temperature. For example, they found that at 350 ℃, the pira (Piria) reaction predominates and produces acetone and carbon dioxide, and they also observed the decomposition products methane and carbon dioxide. However at lower temperatures these decomposition products are reduced. On the other hand, a conventional (normal) reduction reaction leading to the formation of acetaldehyde and ethanol was observed at 300 ℃.

From the foregoing it is evident that existing processes do not have the necessary selectivity to ethylene or that the prior art specifies expensive starting materials other than acetic acid and/or aims to produce products other than ethylene.

Disclosure of Invention

It has now been unexpectedly found that ethylene can be prepared directly from acetic acid on an industrial scale with high selectivity and yield. More specifically, the present invention provides a process for selectively forming ethylene from acetic acid, the process comprising: in a single reaction zone, acetic acid is hydrogenated in the presence of hydrogen over a suitable hydrogenation catalyst, and the intermediates are converted to form ethylene. As examples of such catalysts, the following catalyst metals may be used: copper, cobalt, ruthenium, nickel, aluminum, chromium, zinc, and mixtures thereof.

Detailed Description

The invention is described in detail below with reference to a number of embodiments for illustrative purposes only. Various modifications to the specific embodiments within the spirit and scope of the invention as set forth in the appended claims will be readily apparent to those skilled in the art.

Unless more specifically limited as follows, the terms used herein take their ordinary meaning. Unless otherwise indicated,% and the like refer to mole%.

"conversion" is expressed as a mole percent based on acetic acid in the feed.

"selectivity" is expressed as mole% based on converted acetic acid. For example, if the conversion is 50 mole% and 50 mole% of the converted acetic acid is converted to ethylene, it means that the ethylene selectivity is 50%. The selectivity was calculated from Gas Chromatography (GC) data using the following equation:

without intending to be bound by theory, it is believed that the conversion of acetic acid to ethylene according to the present invention proceeds according to one or more of the following chemical equations:

step 1 a: the acetic acid is hydrogenated to obtain ethylene.

Step 1 b: the ethanol is obtained by hydrogenating acetic acid.

Step 1 c: the acetic acid is hydrogenated to obtain ethyl acetate.

Step 2 a: the ethyl acetate is cracked to obtain ethylene and acetic acid.

And step 2 b: dehydrating the ethanol to obtain the ethylene.

According to the present invention, the conversion of acetic acid to ethylene is carried out in a single reaction zone, which may be, for example, a single fixed bed. The fixed bed may contain a mixture of different catalyst particles or catalyst particles comprising multiple catalysts. Typically, the reaction zone includes at least a hydrogenation catalyst and optionally also a dehydration catalyst and/or a cracking catalyst.

Various hydrogenation catalysts known to those skilled in the art may be used in hydrogenating acetic acid to ethanol in the first step of the process of the present invention. Suitable hydrogenation catalysts are metal catalysts on a suitable support. As mentioned previously, the following catalysts may be mentioned by way of example of such catalysts without any limitation: copper, cobalt, ruthenium, nickel, aluminum, chromium, zinc, palladium, and mixtures thereof. Typically, a single metal, bimetallic or trimetallic catalyst on a suitable support may be used as the hydrogenation catalyst. Copper alone or in combination with aluminum, chromium or zinc is therefore particularly preferred. Similarly, cobalt alone or in combination with ruthenium is preferred. Examples of additional metals that may be used with cobalt as the second or third metal include, without any limitation: platinum, palladium, rhodium, rhenium, iridium, chromium, copper, tin, molybdenum, tungsten, and vanadium.

Various catalyst supports known in the art may be used to support the catalyst of the present invention. Examples of such supports include, without any limitation, zeolites, iron oxides, silica, alumina, titania, zirconia, magnesia, calcium silicate, carbon, graphite, and mixtures thereof. Preferred supports are H-ZSM-5, iron oxide, silica, calcium silicate, carbon or graphite. It is also important to note that the higher the purity of the silica, the more preferred it is as a support in the present invention.

In embodiments of the invention, specific examples of supported hydrogenation catalysts include zeolites, such as H-ZSM-5, iron oxide, silica, alumina, titania, zirconia, magnesia, calcium silicate, carbon, graphite, and mixtures thereof. In particular, as described above, copper supported on iron oxide, copper-aluminum catalyst, cobalt supported on H-ZSM-5, bimetallic catalyst ruthenium-cobalt supported on silica, cobalt supported on carbon are preferable.

Some commercially available catalysts include the following: copper-aluminum catalyst sold by Sud Chemie under the name T-4489; copper-zinc catalysts sold under the names T-2130, T-4427 and T-4492; copper-chromium catalysts sold under the names T-4419 and G-99B; and nickel catalysts sold under the names NiSAT 310, C47-7-04, G-49 and G-69; all sold by Sud Chemie. Particular preference is given to the copper-aluminum catalysts sold under the name T-4489.

The amount of metal supported on the support is not critical in the present invention and can range from about 3% to about 10% by weight. Metal loadings of about 4 wt% to about 6 wt% based on the weight of the support are particularly preferred. Thus, for example, 4 to 6 wt.% copper supported on iron oxide is a particularly preferred catalyst.

The metal impregnation may be performed using any method known in the art. Typically, the support is dried at 120 ℃ and shaped into particles having a size distribution of about 0.2 to 0.4mm prior to impregnation. Optionally, the carrier may be extruded, crushed and sieved to achieve the desired size distribution. Any known method of shaping the carrier material into the desired size distribution may be used.

For supports with low surface area, such as alpha-alumina or iron oxide, an excess of metal solution is added until fully wet or excess liquid impregnation is achieved in order to obtain the desired metal loading.

As noted above, some hydrogenation catalysts are of the bimetallic type. Typically, in such cases, one metal acts as a promoter metal while the other metal is the primary metal. Such as copper, nickel, cobalt and iron, are considered the primary metals used in the preparation of the hydrogenation catalysts of the present invention. The primary metal may be combined with a promoter metal such as tungsten, vanadium, molybdenum, chromium or zinc. It should be noted, however, that sometimes the primary metal may also act as a promoter metal or vice versa. For example, nickel may be used as a promoter metal when iron is used as the primary metal. Similarly, chromium may be used as the primary metal in combination with copper (i.e., Cu — Cr as the primary bimetallic metal), which may be further combined with a promoter metal such as cerium, magnesium, or zinc.

The bimetallic catalyst is typically impregnated in two steps. First, the "promoter" metal is added, followed by the "primary" metal. Each impregnation step is followed by drying and calcination. Bimetallic catalysts may also be prepared by co-impregnation. In the case of trimetallic Cu/Cr containing catalysts as described above, sequential impregnation may be used with initial addition of the "promoter" metal. The second impregnation step may comprise co-impregnation of the two main metals, i.e. Cu and Cr. For example, it can be prepared on SiO by first impregnating cerium nitrate, followed by co-impregnation of copper nitrate and chromium nitrate₂Cu-Cr-Ce. Again, each impregnation is followed by drying and calcination. In most cases, metallic nitric acid can be usedThe salt solution is impregnated. However, various other soluble salts that release metal ions upon calcination may also be used. Examples of other suitable metal salts for impregnation include metal hydroxides, metal oxides, metal acetates, ammonium metal oxides such as ammonium heptamolybdate hexahydrate, metal acids such as perrhenic acid solution, metal oxalates, and the like.

As already mentioned above, in a further aspect of the process of the present invention, any known zeolite can be used as the supported catalyst. Many zeolite catalysts are known in the art, including synthetic as well as natural, all of which can be used as supported catalysts in the present invention. More specifically, any zeolite having a pore size of at least about 0.6nm may be used, and among these zeolites, a catalyst selected from the group consisting of mordenite, ZSM-5, zeolite X and zeolite Y is preferably used.

The preparation of large pore mordenite is described, for example, in U.S. patent No.4,018,514 and in mol.sieves pap.conf., 1967, 78, soc.chem.ind.london, to d.domine and j.quabex.

Zeolite X is described, for example, in U.S. patent No.2,882,244 and zeolite Y is described in U.S. patent No.3,130,007.

Various zeolites and zeolite-type materials are known in the art for catalysis of chemical reactions. For example, U.S. Pat. No.3,702,886 to Argauer discloses a class of synthetic zeolites referred to as "ZSM-5 zeolites" which are effective for the catalysis of various hydrocarbon conversion processes.

Zeolites suitable for the process of the invention may be in the basic form, partially or fully acidified form, or partially dealuminated form.

In another aspect of the process of the present invention, any known dehydration catalyst may be used in the reaction zone of the process of the present invention. Typically, a zeolite catalyst is used as the dehydration catalyst and may support a dehydrogenation catalyst. While any zeolite having a pore size of at least about 0.6nm may be used, preferably used in such zeolites are dehydration catalysts selected from the group consisting of mordenite, ZSM-5, zeolite X and zeolite Y.

The active dehydration catalyst referred to as an "H-ZSM-5" or "H-mordenite" zeolite in the process of the present invention is prepared from the corresponding "ZSM-5" or "mordenite" zeolite by replacing a substantial portion, usually at least about 80% of the cations of the latter zeolite with hydrogen ions using techniques well known in the art. For example, H-mordenite zeolite is prepared by calcining ammonium form mordenite at 500-550 deg.C for 4-8 hours. If the sodium form of mordenite is used as a precursor, the sodium mordenite is ion-exchanged to the ammonium form prior to calcination.

These zeolite catalysts are essentially crystalline aluminosilicates or a combination of silica and alumina in a neutral form, i.e. in a well-defined crystalline structure. Among the zeolite catalysts of the particularly preferred class for the purposes of the present invention are those in which SiO is present₂With Al₂O₃Is about 10 to 60.

As previously described, ethylene is produced by dehydration and decomposition or "cracking" of ethyl acetate to ethylene and acetic acid. This may be simply done at elevated temperatures due to thermal cracking or may be a catalytic reaction using a cracking catalyst if desired. Suitable cracking catalysts include sulfonic acid resins such as the perfluorosulfonic acid resins disclosed in the above-mentioned U.S. patent No.4,399,305 (the disclosure of which is incorporated herein by reference). Zeolites are also suitable as cracking catalysts, as described in U.S. Pat. No.4,620,050, the disclosure of which is also incorporated herein by reference. Thus, in the highly efficient process of the present invention, a zeolite catalyst can be used to simultaneously dehydrate ethanol to ethylene and decompose ethyl acetate to ethylene.

The selectivity of acetic acid to ethylene in typical cases is suitably greater than 10%, more suitably for example at least 20%, or at least 25% or up to about 40%. Depending on the mixture of by-products, if for undesired products such as CO₂Is kept low, it may be desirable to operate at moderate selectivity and recycle products such as acetaldehyde for further hydrogenation and de-hydrogenationAnd (3) water.

Preferably, for the purposes of the process of the present invention, suitable hydrogenation catalysts are copper on iron oxide or copper-aluminum catalysts sold under the trade name T-4489 by Sud Chemie, cobalt supported on H-ZSM-5, bimetallic catalysts, ruthenium and cobalt supported on silica, and cobalt supported on carbon. In this embodiment of the process of the present invention, the copper loading on the iron oxide support or in the bimetallic copper-aluminum catalyst is typically from about 3 wt.% to about 10 wt.%, preferably it is from about 4 wt.% to about 6 wt.%. Similarly, the cobalt loading on H-ZSM-5 or silica or carbon is typically about 5 wt%. The amount of ruthenium in the bimetallic catalyst is also about 5 wt.%.

In another aspect of the process of the invention, the hydrogenation and dehydration of acetic acid is carried out at a pressure just sufficient to overcome the pressure drop across the catalytic bed.

The reaction can be carried out under a wide variety of conditions, in either the gaseous or liquid state. Preferably, the reaction is carried out in the gas phase. Reaction temperatures of, for example, from about 200 ℃ to about 375 ℃, preferably from about 250 ℃ to about 350 ℃, may be used. The pressure is generally not critical to the reaction and subatmospheric, atmospheric or superatmospheric pressures may be used. However, in most cases, the reaction pressure may be about 1 to 30 atmospheres absolute.

Although the reaction consumes 2 moles of hydrogen per mole of acetic acid to produce 1 mole of ethylene, the actual molar ratio of acetic acid to hydrogen in the feed stream may vary over a wide range, e.g., from about 100: 1 to 1: 100. However, it is preferred that such a ratio is from about 1: 20 to 1: 2.

The feed used in connection with the process of the present invention may be obtained from any suitable source, including natural gas, petroleum, coal, biomass, and the like. It is well known to produce acetic acid by methanol carbonylation, acetaldehyde oxidation, ethylene oxidation, oxidative fermentation, anaerobic fermentation, and the like. As petroleum and natural gas become more expensive, processes for producing acetic acid and intermediates such as methanol and carbon monoxide from alternative carbon sources have attracted greater attention. Of particular interest is the production of acetic acid from synthesis gas (syngas) which may be obtained from any suitable carbon source. U.S. patent No.6,232,352 to Vidalin, the disclosure of which is incorporated herein by reference, for example, teaches a method of retrofitting a methanol plant for the manufacture of acetic acid. By retrofitting a methanol plant, the substantial capital costs associated with CO production for a new acetic acid plant are significantly reduced or largely eliminated. All or a portion of the syngas is split from the methanol synthesis loop (loop) and fed to a separation unit to recover CO and hydrogen, which are then used to produce acetic acid. In addition to acetic acid, the process may also be used to produce hydrogen for use in connection with the present invention.

U.S. patent No. re 35,377 to Steinberg et al, also incorporated herein by reference, provides a process for the production of methanol by the conversion of carbonaceous materials such as oil, coal, natural gas and biomass materials. The method comprises hydro-gasification of solid and/or liquid carbonaceous material to obtain a process gas, which is steam pyrolyzed together with additional natural gas to form synthesis gas. The synthesis gas is converted to methanol which can be carbonylated to acetic acid. This process also produces hydrogen for use in connection with the present invention as described above. See also U.S. Pat. No.5,821,111 to Grady et al, which discloses a process for converting waste biomass to syngas by gasification, and U.S. Pat. No.6,685,754 to Kindig et al, the disclosures of which are incorporated herein by reference.

The acetic acid may be vaporized at the reaction temperature and then may be fed with hydrogen either undiluted or diluted with a relatively inert carrier gas such as nitrogen, argon, helium, carbon dioxide, and the like.

Alternatively, acetic acid in vapor form may be withdrawn directly as a crude product from the flash vessel of a methanol carbonylation unit of the type described in U.S. Pat. No.6,657,078 to Scates et al, the disclosure of which is incorporated herein by reference. The crude vapor product can be fed directly to the reaction zone of the present invention without the need to condense acetic acid and light ends or remove water, thereby saving overall processing costs.

The contact or residence time may also vary widely depending on such variables as the amount of acetic acid, catalyst, reactor, temperature and pressure. Typical contact times range from fractions of a second to more than several hours when using catalyst systems other than fixed beds, with contact times of about 0.5 to 100 seconds being preferred, at least for gas phase reactions.

Typically, the catalyst is used in a fixed bed reactor, for example in the shape of an elongated tube or pipe, in which reactants, typically in vapor form, pass over or through the catalyst. Other reactors, such as fluidized bed or ebullated bed reactors, can also be used if desired. In some cases it is advantageous to use a catalyst bed in combination with an inert material such as glass wool to regulate the pressure drop of the reactant stream across the catalyst bed and the contact time of the reactant compounds with the catalyst particles.

Also provided in a preferred embodiment is a process for selectively forming ethylene from acetic acid, the process comprising: contacting a feed stream of acetic acid and hydrogen with a catalyst selected from copper supported on iron oxide, copper-aluminum catalyst, cobalt supported on H-ZSM-5, ruthenium-cobalt supported on silica or cobalt supported on carbon at a temperature of about 250 ℃ to 350 ℃ to form ethylene.

In this embodiment of the process of the invention, the preferred catalyst is 5 wt% copper on iron oxide, 5 wt% cobalt on H-ZSM-5, 5 wt% cobalt and 5 wt% ruthenium on silica or 5 wt% cobalt on carbon. In this embodiment of the process of the invention the reaction is preferably carried out in the gas phase in a tubular reactor packed with a catalyst bed and at a temperature of about 250 ℃ to 350 ℃ and a pressure of about 1 to 30 atmospheres absolute, the contact time of the reactants being about 0.5 to 100 seconds.

The following examples describe the procedures used to prepare the various catalysts used in the process of the present invention.

Example A

Preparation of 5 wt.% copper on iron oxide

Powdered and sieved iron oxide (100g) with a uniform particle size distribution of about 0.2mm was dried overnight at 120 ℃ in an oven under nitrogen atmosphere and then cooled to room temperature. To this was added a solution (100ml) of copper nitrate pentahydrate (17g) in distilled water. The resulting slurry was dried in an oven with gradual heating to 110 ℃ (> 2 hours, 10 ℃/min). The impregnated catalyst mixture was then calcined at 500 deg.C (6 hours, 1 deg.C/min).

Example B

Preparation of 5 wt% cobalt on H-ZSM-5

Example a was substantially repeated except that a suitable amount of cobalt nitrate hexahydrate as the metal salt and H-ZSM-5 as the supported catalyst were used to prepare 5 wt% cobalt supported on H-ZSM-5.

Example C

Preparation of 5 wt% cobalt and 5 wt% ruthenium on silica

Example a was substantially repeated except that suitable amounts of cobalt nitrate hexahydrate and ruthenium nitrosyl nitrate were used as the metal salts and silica as the supported catalyst to produce 5 wt% cobalt and 5 wt% ruthenium supported on silica.

Example D

Preparation of 5 wt% cobalt on carbon

Example a was substantially repeated except that a suitable amount of cobalt nitrate hexahydrate as the metal salt and carbon as the supported catalyst were used to prepare 5 wt% cobalt supported on carbon.

Gas Chromatography (GC) analysis of the product

Product analysis was performed by online GC. The reactants and products were analyzed using a three-channel integrated GC equipped with 1 Flame Ionization Detector (FID) and 2 Thermal Conductivity Detectors (TCD). The front channel was equipped with FID and CP-Sil 5(20m) + WaxFFap (5m) columns and used for quantitation:

acetaldehyde

Ethanol

Acetone (II)

Acetic acid methyl ester

Vinyl acetate (VAA)

Ethyl acetate

Acetic acid

Ethylene diacetate

Ethylene glycol

Ethylidene diacetate

Paraldehyde

The middle channel was equipped with TCD and Porabond Q columns and was used for quantitation:

CO₂

ethylene

Ethane (III)

The back channel was equipped with TCD and Molsieve 5A columns and used for quantitation:

helium gas

Hydrogen gas

Nitrogen gas

Methane

Carbon monoxide

Prior to the reaction, the retention times of the different components were determined by spiking with individual compounds and the GC was calibrated with a calibration gas of known composition or with a liquid solution of known composition. This allows the response factors of the individual components to be determined.

Example 1

The catalyst used was 5 wt% copper on iron oxide prepared according to the procedure of example a.

In a tubular reactor made of stainless steel with an internal diameter of 30mm and capable of being brought to a controlled temperature, 50ml of a catalyst of 5% by weight of copper on iron oxide were placed. The length of the catalyst bed after loading was approximately about 70 mm.

The feed liquid consists essentially of acetic acid. The reaction feed liquid was vaporized and charged to the reactor at an average total Gas Hourly Space Velocity (GHSV) of about 2500h "1 along with hydrogen and helium as a carrier gas at a temperature of about 350 ℃ and a pressure of 100 psig. The resulting feed stream contains a mole percent of acetic acid of from about 4.4% to about 13.8% and a mole percent of hydrogen of from about 14% to about 77%. A portion of the vapor effluent was passed through a gas chromatograph for analysis of the effluent contents. The results are shown in table 1. The selectivity to ethylene was 16% at 100% conversion of acetic acid.

Example 2

The catalyst used was 5 wt% cobalt on H-ZSM-5 prepared according to the procedure of example B.

The procedure given in example 1 was substantially repeated, the mean total Gas Hourly Space Velocity (GHSV) of the feed stream of acetic acid, hydrogen and helium vaporized at a temperature of 250 ℃ and a pressure of 1 bar being 10,000h^-1. A portion of the vapor effluent was passed through a gas chromatograph for analysis of the effluent contents. The results are shown in table 1. The acetic acid conversion was 3% and the ethylene selectivity was 28%.

Example 3

The catalyst used was a bimetallic catalyst containing 5% by weight of cobalt and 5% by weight of ruthenium supported on silica prepared according to the procedure of example C.

The procedure given in example 1 was substantially repeated, the average total Gas Hourly Space Velocity (GHSV) of the feed stream of acetic acid, hydrogen and helium vaporized at a temperature of 350 ℃ and a pressure of 1 bar being 2500h "1. A portion of the vapor effluent was passed through a gas chromatograph for analysis of the effluent contents. The results are shown in table 1. The acetic acid conversion was 4% and the ethylene selectivity was 14%.

Example 4

The catalyst used was 5 wt% cobalt supported on carbon prepared according to the procedure of example D.

The procedure given in example 1 was substantially repeated, the average total Gas Hourly Space Velocity (GHSV) of the feed stream of acetic acid, hydrogen and helium vaporized at a temperature of 350 ℃ and a pressure of 1 bar being 2500h "1. A portion of the vapor effluent was passed through a gas chromatograph for analysis of the effluent contents. The results are shown in table 1. The acetic acid conversion was 2% and the ethylene selectivity was 12%.

Generally, selectivities to ethylene of greater than about 10% are highly desirable; it is appreciated that other byproducts, such as ethanol or ethyl acetate, may be recycled to the reactor along with unreacted acetic acid, although other byproducts may also be reprocessed or used for fuel value. Expected to be on CO₂Less than 10%, preferably less than 5%.

TABLE 1 acetic acid conversion and Selectivity

Examples	Ethylene selectivity (%)	Acetic acid conversion (%)	Other products
				1	16	100	Acetaldehyde-31%, ethane-15%, ethyl acetate-4%, CO2-5％
2	29	3	Acetaldehyde-51%, ethane-28%,
				3	14	4	acetaldehyde-78%, ethane-8%
4	12	2	Acetaldehyde-8%, methane-47%, ethane-5%

Comparative examples 1 to 5

These examples illustrate the reaction of acetic acid and hydrogen over various catalysts, wherein no ethylene is formed and/or very low ethylene levels are detected.

In all of these examples, the procedure given in example 1 was followed essentially except that the different catalysts listed in Table 2 were used. The reaction temperature and selectivity to ethylene are also listed in table 2.

TABLE 2

Various other products including acetaldehyde, ethanol, ethyl acetate, ethane, carbon monoxide, carbon dioxide, methane, isopropanol, acetone, and water were observed in these examples.

While the invention has been described by certain of the foregoing embodiments, it is not to be construed as being limited thereby; but that the invention include the general scope disclosed above. Various modifications and implementations can be made without departing from the spirit and scope thereof.

Claims (25)

1. A process for the selective and direct formation of ethylene from acetic acid comprising: a feed stream containing acetic acid and hydrogen is contacted with a suitable hydrogenation catalyst at an elevated temperature in a single reaction zone, optionally including a dehydration catalyst or a cracking catalyst, to form ethylene.

2. The process according to claim 1, wherein the hydrogenation is carried out on a supported hydrogenation catalyst selected from the group consisting of copper, cobalt, ruthenium, nickel, aluminum, chromium, zinc, palladium and mixtures thereof.

3. The process according to claim 2, wherein the support is selected from the group consisting of iron oxide, H-ZSM-5, silica, alumina, titania, zirconia, magnesia, calcium silicate, carbon, graphite and mixtures thereof.

4. The process according to claim 2, wherein the hydrogenation catalyst is selected from the group consisting of copper supported on iron oxide, copper-aluminum catalyst, copper-zinc catalyst, copper-chromium catalyst, cobalt supported on H-ZSM-5, ruthenium-cobalt supported on silica, cobalt supported on carbon, and nickel catalyst.

5. The process according to claim 2, wherein the hydrogenation catalyst is selected from the group consisting of copper supported on iron oxide, copper-aluminum catalyst, cobalt supported on H-ZSM-5, ruthenium-cobalt supported on silica or cobalt supported on carbon.

6. The process according to claim 1, wherein the hydrogenation catalyst is copper supported on iron oxide, cobalt supported on H-ZSM-5, ruthenium-cobalt supported on silica or cobalt supported on carbon.

7. The process according to claim 6, wherein the catalyst is copper supported on iron oxide.

8. The process according to claim 6, wherein the catalyst is cobalt supported on H-ZSM-5.

9. The process according to claim 6, wherein the catalyst is ruthenium-cobalt supported on silica or cobalt supported on carbon.

10. The process according to claim 6, wherein the loading of copper on iron oxide is from about 3% to about 10% by weight.

11. The process according to claim 7, wherein the loading of copper on iron oxide is from about 4% to about 6% by weight.

12. The process according to claim 6, wherein the loading of cobalt on H-ZSM-5, silica or carbon is in the range of about 3 wt% to about 10 wt%.

13. The process according to claim 12, wherein the loading of cobalt on H-ZSM-5, silica or carbon is in the range of from about 4 wt% to about 6 wt%.

14. The process according to claim 6 wherein the loading of ruthenium on silica is from about 3% to about 10% by weight.

15. The process according to claim 6, wherein the loading of ruthenium on silica is from about 4 wt% to about 6 wt%.

16. The process according to claim 1, wherein the hydrogenation is carried out at a pressure just sufficient to overcome the pressure drop across the catalytic bed.

17. The process according to claim 1, wherein the hydrogenation is carried out in the gas phase and at a temperature of about 200 ℃ to 375 ℃.

18. The process according to claim 17, wherein the hydrogenation is carried out in the gas phase and at a temperature of about 250 ℃ to 350 ℃.

19. The process according to claim 17 wherein the catalyst is in the form of a layered fixed bed and the feed stream to the bed further contains an inert carrier gas.

20. The process of claim 17 wherein the reactants are comprised of acetic acid and hydrogen in a molar ratio of from about 100: 1 to about 1: 100, the reaction temperature is from about 250 ℃ to about 350 ℃, the reaction pressure is from about 1 to about 30 atmospheres absolute, and the contact time of the reactants and catalyst is from about 0.5 to about 100 seconds.

21. The process of claim 17 wherein the reactants are comprised of acetic acid and hydrogen in a molar ratio of from about 1: 20 to about 1: 2, the reaction temperature is from about 300 ℃ to about 350 ℃, the reaction pressure is from about 1 to about 30 atmospheres absolute, and the contact time of the reactants and catalyst is from about 0.5 to about 100 seconds.

22. A process for selectively forming ethylene from acetic acid, the process comprising: contacting a feed stream of acetic acid and hydrogen with a hydrogenation catalyst selected from copper supported on iron oxide, a copper-aluminum catalyst, cobalt supported on H-ZSM-5, ruthenium-cobalt supported on silica or cobalt supported on carbon at a temperature of about 250 ℃ to 350 ℃ to form ethylene.

23. The process according to claim 22, wherein the hydrogenation catalyst is 5 wt% copper on iron oxide.

24. The process according to claim 22, wherein the hydrogenation catalyst is 5 wt% cobalt supported on H-ZSM-5.

25. The process according to claim 22, wherein the hydrogenation and dehydration catalyst is layered in a fixed bed and the reaction is carried out in the vapor phase at a temperature of about 300 ℃ to 350 ℃ and at an absolute atmospheric pressure of about 1 to 30, the contact time of the reactants being about 0.5 to 100 seconds.