Process for removing water from alcohol mixtures

Abstract

A process for removing water from an alcohol mixture, such as an ethanol mixture in the production of ethanol by hydrogenating acetic acid. The water is removed from a distillate of a first column using an adsorption unit, membrane, extractive column distillation, or a combination thereof.

Description

FIELD OF THE INVENTION

The present invention relates generally to processes for producing alcohol and, in particular, to a process for removing water from an ethanol mixture that comprises ethanol, water and ethyl acetate.

BACKGROUND OF THE INVENTION

Ethanol for industrial use is conventionally produced from petrochemical feed stocks, such as oil, natural gas, or coal, from feed stock intermediates, such as syngas, or from starchy materials or cellulose materials, such as corn or sugar cane. Conventional methods for producing ethanol from petrochemical feed stocks, as well as from cellulose materials, include the acid-catalyzed hydration of ethylene, methanol homologation, direct alcohol synthesis, and Fischer-Tropsch synthesis. Instability in petrochemical feed stock prices contributes to fluctuations in the cost of conventionally produced ethanol, making the need for alternative sources of ethanol production all the greater when feed stock prices rise. Starchy materials, as well as cellulose material, are converted to ethanol by fermentation. However, fermentation is typically used for consumer production of ethanol, which is suitable for fuels or human consumption. In addition, fermentation of starchy or cellulose materials competes with food sources and places restraints on the amount of ethanol that can be produced for industrial use.

Ethanol production via the reduction of alkanoic acids and/or other carbonyl group-containing compounds has been widely studied, and a variety of combinations of catalysts, supports, and operating conditions have been mentioned in the literature. During the reduction of alkanoic acid, e.g., acetic acid, other compounds are formed with ethanol or are formed in side reactions. These impurities limit the production and recovery of ethanol from such reaction mixtures. For example, during hydrogenation, esters are produced that together with ethanol and/or water form azeotropes, which are difficult to separate. In addition when conversion is incomplete, unreacted acetic acid remains in the crude ethanol product, which must be removed to recover ethanol.

EP02060553 describes a process for converting hydrocarbons to ethanol involving converting the hydrocarbons to ethanoic acid and hydrogenating the ethanoic acid to ethanol. The stream from the hydrogenation reactor is separated to obtain an ethanol stream and a stream of acetic acid and ethyl acetate, which is recycled to the hydrogenation reactor.

The need remains for improved processes for recovering ethanol from a crude product obtained by reducing alkanoic acids, such as acetic acid, and/or other carbonyl group-containing compounds.

SUMMARY OF THE INVENTION

In a first embodiment, the present invention is directed to a process for producing ethanol, comprising hydrogenating acetic acid in a reactor in the presence of a catalyst to form a crude ethanol product; separating a portion of the crude ethanol product in a first distillation column to yield a first residue comprising acetic acid and a first distillate comprising ethanol, ethyl acetate, and water; removing water from a portion of the first distillate to yield an ethanol mixture stream comprising less than 10 wt. % water; and recovering ethanol from the ethanol mixture stream.

In a second embodiment, the present is directed to a process for producing ethanol, comprising hydrogenating acetic acid in a reactor in the presence of a catalyst to form a crude ethanol product; separating a portion of the crude ethanol product in a first distillation column to yield a first residue comprising acetic acid and a first distillate comprising ethanol, ethyl acetate, and water; removing water from the first distillate to yield an ethanol mixture stream; and separating a portion of the ethanol mixture stream in a second distillation column to yield a second residue comprising ethanol and a second distillate comprising ethyl acetate.

In a third embodiment, the present invention is directed to a method for recovering ethanol, comprising providing a crude ethanol product comprising ethanol, acetic acid, ethyl acetate, and water; separating a portion of the crude ethanol product in a first distillation column to yield a first residue comprising a majority of the acetic acid and a first distillate comprising a majority of each of the ethanol, and the ethyl acetate; removing water from the first distillate to yield an ethanol mixture stream comprising less than 10 wt. % water; and recovering ethanol from the ethanol mixture stream.

BRIEF DESCRIPTION OF DRAWINGS

The invention may be more completely understood in consideration of the following detailed description of various embodiments of the invention in connection with the accompanying drawings, wherein like numerals designate similar parts.

FIG. 1 is a schematic diagram of an ethanol production system with a water separator for removing water from the first distillate in accordance with one embodiment of the present invention.

FIG. 2 is a schematic diagram of an ethanol production system with a membrane for removing water from the first distillate in accordance with one embodiment of the present invention.

FIG. 3 is a schematic diagram of a bypass line for a water separator for removing water from the first distillate in accordance with one embodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The present invention relates to processes for recovering ethanol produced by hydrogenating acetic acid in the presence of a catalyst. The hydrogenation reaction produces a crude ethanol product that comprises ethanol, water, ethyl acetate, unreacted acetic acid, and other impurities. The concentration of some of these compounds in the reaction mixture is largely a factor of catalyst composition and process conditions. However, water concentration in the reaction mixture is not dictated by these factors because water is co-produced with ethanol in the hydrogenation reaction in about a 1:1 molar ratio. Thus, producing additional ethanol also results in the production of additional water.

In one embodiment of the present invention, the process involves introducing the crude ethanol product to an initial separation column (first column), which separates the crude ethanol product into a distillate comprising ethanol, ethyl acetate and water, and a residue comprising water and unreacted acetic acid. Water is then removed from the distillate to form an ethanol mixture stream, preferably comprising less than 10 wt. % water, less than 6 wt. % water or less than 4 wt. % water. In terms of ranges the ethanol mixture stream comprises from 0.001 to 10 wt. % water, e.g., from 0.01 to 6 wt. % water or from 0.1 to 4 wt. % water. Product ethanol is then recovered from the ethanol mixture stream. It has now been discovered that removing water in the distillate of the initial column may improve overall separation efficiency in recovering ethanol.

Water and ethanol form an azeotrope that is difficult to separate in a distillation column. The ethanol-water azeotrope limits the recoverable ethanol in distillation columns to an ethanol product comprising about 92-96 wt. % of ethanol. The energy required to approach this azeotrope in a distillation column, regardless of the presence of other compounds, is significant. The present invention involves using less energy in the initial column than would be required to approach the azeotrope, resulting in water being carried overhead in the distillate. Water is then removed from the distillate using a water separator, which beneficially requires less energy than is required for approaching the water/ethanol azeotrope in a distillation column. Thus, the present invention provides a low energy approach for dehydrating a crude ethanol product and thus removing water that is co-produced with ethanol.

The concentration of water in the distillate may vary depending on the acetic acid conversion. In one embodiment, the distillate comprises water in an amount greater than the amount of water in the ethanol/water azeotrope, e.g., in an amount greater than 4 wt. %, greater than 5 wt. %, or greater than 7 wt. %. In terms of ranges, the distillate optionally comprises water in an amount from 4 wt. % to 38 wt. %, e.g., from 7 wt. % to 32 wt. %, or from 7 wt. % to 25 wt. %.

Because the water concentration in the distillate is typically greater than the acceptable amount of water for industrial or fuel grade ethanol applications, in one embodiment of the present invention, the process involves removing a substantial portion of the water from the distillate to produce an ethanol mixture. Preferably, the water is removed before separating any appreciable amount of organics, ethyl acetate or acetaldehyde. In one embodiment, the water is removed prior to condensing the distillate. For example, distillate in vapor phase may be fed to an adsorption unit comprising a molecular sieve or a membrane. In some embodiments, distillate may be condensed to a liquid and fed to a membrane. The heat of vaporization for water is provided to the distillate to allow water to permeate through the membrane. In preferred embodiments, at least 50% of the water in the distillate is removed, e.g., at least 60% of the water or at least 75% of the water, based on the total amount of water in the distillate. In more preferred embodiments, from 90 to 99% of the water may be removed from the distillate. Thus, the resulting ethanol mixture may comprise only a minor amount of water, from 0.01 to 10 wt. %, e.g., from 0.5 to 6 wt. %, or from 0.5 to 4 wt. %. In one embodiment, the ethanol mixture comprises a water concentration that is less than the amount of water in the ethanol/water azeotrope. In order to achieve a water concentration that is below the amount of water in the ethanol/water azeotrope, a large of amount of energy is required. Thus, the present invention beneficially removes water from the first distillate to yield an ethanol mixture without a large amount of energy. Also, because the ethanol mixture comprises less water, the need to remove water during the later stage of product separation is also reduced.

In an exemplary embodiment, the energy requirements by the initial column in the process according to the present invention may be less than 5.5 MMBtu per ton of refined ethanol, e.g., less than 4.5 MMBtu per ton of refined ethanol or less than 3.5 MMBtu per ton of refined ethanol. In some embodiments, the process may operate with higher energy requirements provided that the total energy requirement is less than the energy required to remove most of the water from the crude ethanol product in the distillate, e.g. more than 65% of the water in the crude ethanol product.

The water that is removed from the distillate may be returned to the initial column and ultimately removed from the initial column via the residue. In one embodiment, a portion of the removed water may be condensed and returned below the feed point of the crude ethanol product to the initial column, e.g., near the bottom of the initial column. Depending on the water removal technique, there may be some ethanol and ethyl acetate in the removed water and thus it may be desirable to recover these compounds by returning at least a portion of the removed water to the initial column. Returning the removed water to the initial column may increase the amount of water withdrawn as the residue. In other embodiments, a portion of the removed water may be fed to a separation column, e.g. light-ends column, used in recovering an ethanol product from the ethanol mixture. The presence of a small amount of water, e.g., less than 10 wt. % water based on the total feed, in the light-ends column may be beneficial in facilitating the separation of ethyl acetate from ethanol. A portion of the removed water may also be purged as needed to remove water from the system.

The ethanol mixture may be further processed in the light-ends column to recover ethanol. In some embodiments, it may be desirable to maintain a concentration of water in the light-ends column. Depending on the type of water separator, the ethanol mixture may comprises less than 0.5 wt. % water. To control the water concentration, a by-pass line may be used to split the distillate. The split ratio may vary to control the amount of water in feed to the light-ends column. In one embodiment, the split ratio may range from 10:1 to 1:10, e.g., from 5:1 to 1:5 or about 1:1. Other split ratios may be used when controlling the water concentration. The distillate in the by-pass line is not separated to remove water and may be combined or co-fed with the ethanol mixture to the light-ends column. The combined distillate and ethanol mixture has a total water concentration of greater than 0.5 wt. %, e.g., greater than 2 wt. % or greater than 5 wt. %. In terms of ranges, the total water concentration of the combined distillate and ethanol mixture may be from 0.5 to 15 wt. %, e.g., from 2 to 12 wt. %, or from 5 to 10 wt. %. The additional water for the light-ends column is typically recovered with the ethanol and separated as desired to provide an ethanol product.

The process of the present invention may use any suitable technique for removing water from the distillate. For example, water may be removed in the vapor phase, before condensation, or in the liquid phase. Water may be removed, for example, using an adsorption unit, membrane, molecular sieves, extractive column distillation, or a combination thereof. Suitable adsorption units include pressure swing adsorption (PSA) units and thermal swing adsorption (TSA) units. The adsorption units may comprises molecular sieves, such as aluminosilicate compounds.

A membrane or an array of membranes may also be employed to separate water from the distillate. The membrane or array of membranes may be selected from any suitable membrane that is capable of removing a permeate water stream from a stream that also comprises ethanol and ethyl acetate.

Hydrogenation of Acetic Acid

The process of the present invention may be used with any hydrogenation process for producing ethanol. The materials, catalysts, reaction conditions, and separation processes that may be used in the hydrogenation of acetic acid are described further below.

The raw materials, acetic acid and hydrogen, used in connection with the process of this invention may be derived from any suitable source including natural gas, petroleum, coal, biomass, and so forth. As examples, acetic acid may be produced via methanol carbonylation, acetaldehyde oxidation, ethylene oxidation, oxidative fermentation, and anaerobic fermentation. Methanol carbonylation processes suitable for production of acetic acid are described in U.S. Pat. Nos. 7,208,624; 7,115,772; 7,005,541; 6,657,078; 6,627,770; 6,143,930; 5,599,976; 5,144,068; 5,026,908; 5,001,259; and 4,994,608, the entire disclosures of which are incorporated herein by reference. Optionally, the production of ethanol may be integrated with such methanol carbonylation processes.

As petroleum and natural gas prices fluctuate becoming either more or less expensive, methods for producing acetic acid and intermediates such as methanol and carbon monoxide from alternate carbon sources have drawn increasing interest. In particular, when petroleum is relatively expensive, it may become advantageous to produce acetic acid from synthesis gas (“syngas”) that is derived from more available carbon sources. U.S. Pat. No. 6,232,352, the entirety 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 large capital costs associated with CO generation for a new acetic acid plant are significantly reduced or largely eliminated. All or part of the syngas is diverted from the methanol synthesis loop and supplied to a separator unit to recover CO, which are then used to produce acetic acid. In a similar manner, hydrogen for the hydrogenation step may be supplied from syngas.

In some embodiments, some or all of the raw materials for the above-described acetic acid hydrogenation process may be derived partially or entirely from syngas. For example, the acetic acid may be formed from methanol and carbon monoxide, both of which may be derived from syngas. The syngas may be formed by partial oxidation reforming or steam reforming, and the carbon monoxide may be separated from syngas. Similarly, hydrogen that is used in the step of hydrogenating the acetic acid to form the crude ethanol product may be separated from syngas. The syngas, in turn, may be derived from variety of carbon sources. The carbon source, for example, may be selected from the group consisting of natural gas, oil, petroleum, coal, biomass, and combinations thereof. Syngas or hydrogen may also be obtained from bio-derived methane gas, such as bio-derived methane gas produced by landfills or agricultural waste.

In another embodiment, the acetic acid used in the hydrogenation step may be formed from the fermentation of biomass. The fermentation process preferably utilizes an acetogenic process or a homoacetogenic microorganism to ferment sugars to acetic acid producing little, if any, carbon dioxide as a by-product. The carbon efficiency for the fermentation process preferably is greater than 70%, greater than 80% or greater than 90% as compared to conventional yeast processing, which typically has a carbon efficiency of about 67%. Optionally, the microorganism employed in the fermentation process is of a genus selected from the group consisting of Clostridium, Lactobacillus, Moorella, Thermoanaerobacter, Propionibacterium, Propionispera, Anaerobiospirillum, and Bacteriodes, and in particular, species selected from the group consisting of Clostridium formicoaceticum, Clostridium butyricum, Moorella thermoacetica, Thermoanaerobacter kivui, Lactobacillus delbrukii, Propionibacterium acidipropionici, Propionispera arboris, Anaerobiospirillum succinicproducens, Bacteriodes amylophilus and Bacteriodes ruminicola. Optionally in this process, all or a portion of the unfermented residue from the biomass, e.g., lignans, may be gasified to form hydrogen that may be used in the hydrogenation step of the present invention. Exemplary fermentation processes for forming acetic acid are disclosed in U.S. Pat. Nos. 6,509,180; 6,927,048; 7,074,603; 7,507,562; 7,351,559; 7,601,865; 7,682,812; and 7,888,082, the entireties of which are incorporated herein by reference. See also U.S. Pub. Nos. 2008/0193989 and 2009/0281354, the entireties of which are incorporated herein by reference.

Examples of biomass include, but are not limited to, agricultural wastes, forest products, grasses, and other cellulosic material, timber harvesting residues, softwood chips, hardwood chips, tree branches, tree stumps, leaves, bark, sawdust, off-spec paper pulp, corn, corn stover, wheat straw, rice straw, sugarcane bagasse, switchgrass, miscanthus, animal manure, municipal garbage, municipal sewage, commercial waste, grape pumice, almond shells, pecan shells, coconut shells, coffee grounds, grass pellets, hay pellets, wood pellets, cardboard, paper, plastic, and cloth. See, e.g., U.S. Pat. No. 7,884,253, the entirety of which is incorporated herein by reference. Another biomass source is black liquor, a thick, dark liquid that is a byproduct of the Kraft process for transforming wood into pulp, which is then dried to make paper. Black liquor is an aqueous solution of lignin residues, hemicellulose, and inorganic chemicals.

U.S. Pat. No. RE 35,377, also incorporated herein by reference, provides a method for the production of methanol by conversion of carbonaceous materials such as oil, coal, natural gas and biomass materials. The process includes hydrogasification of solid and/or liquid carbonaceous materials to obtain a process gas which is steam pyrolized with additional natural gas to form synthesis gas. The syngas is converted to methanol which may be carbonylated to acetic acid. The method likewise produces hydrogen which may be used in the hydrogenation step of this invention as noted above. U.S. Pat. No. 5,821,111, which discloses a process for converting waste biomass through gasification into synthesis gas, and U.S. Pat. No. 6,685,754, which discloses a method for the production of a hydrogen-containing gas composition, such as a synthesis gas including hydrogen and carbon monoxide, are incorporated herein by reference in their entireties.

The acetic acid fed to the hydrogenation reaction may also comprise other carboxylic acids and anhydrides, as well as acetaldehyde and acetone. Preferably, a suitable acetic acid feed stream comprises one or more of the compounds selected from the group consisting of acetic acid, acetic anhydride, acetaldehyde, ethyl acetate, and mixtures thereof. These other compounds may also be hydrogenated in the processes of the present invention. In some embodiments, the presence of carboxylic acids, such as propanoic acid or its anhydride, may be beneficial in producing propanol. Water may also be present in the acetic acid feed.

Alternatively, acetic acid in vapor form may be taken directly as crude product from the flash vessel of a methanol carbonylation unit of the class described in U.S. Pat. No. 6,657,078, the entirety of which is incorporated herein by reference. The crude vapor product, for example, may be fed directly to the ethanol synthesis reaction zones of the present invention without the need for condensing the acetic acid and light ends or removing water, saving overall processing costs.

The acetic acid may be vaporized at the reaction temperature, following which the vaporized acetic acid may be fed along with hydrogen in an undiluted state or diluted with a relatively inert carrier gas, such as nitrogen, argon, helium, carbon dioxide and the like. For reactions run in the vapor phase, the temperature should be controlled in the system such that it does not fall below the dew point of acetic acid. In one embodiment, the acetic acid may be vaporized at the boiling point of acetic acid at the particular pressure, and then the vaporized acetic acid may be further heated to the reactor inlet temperature. In another embodiment, the acetic acid mixed with other gases before vaporizing, followed by heating the mixed vapors up to the reactor inlet temperature. Preferably, the acetic acid is transferred to the vapor state by passing hydrogen and/or recycle gas through the acetic acid at a temperature at or below 125° C., followed by heating of the combined gaseous stream to the reactor inlet temperature.

Some embodiments of the process of hydrogenating acetic acid to form ethanol may include a variety of configurations using a fixed bed reactor or a fluidized bed reactor. In many embodiments of the present invention, an “adiabatic” reactor can be used; that is, there is little or no need for internal plumbing through the reaction zone to add or remove heat. In other embodiments, a radial flow reactor or reactors may be employed, or a series of reactors may be employed with or with out heat exchange, quenching, or introduction of additional feed material. Alternatively, a shell and tube reactor provided with a heat transfer medium may be used. In many cases, the reaction zone may be housed in a single vessel or in a series of vessels with heat exchangers therebetween.

In preferred embodiments, the catalyst is employed in a fixed bed reactor, e.g., in the shape of a pipe or tube, where the reactants, typically in the vapor form, are passed over or through the catalyst. Other reactors, such as fluid or ebullient bed reactors, can be employed. In some instances, the hydrogenation catalysts may be used in conjunction with an inert material to regulate the pressure drop of the reactant stream through the catalyst bed and the contact time of the reactant compounds with the catalyst particles.

The hydrogenation reaction may be carried out in either the liquid phase or vapor phase. Preferably, the reaction is carried out in the vapor phase under the following conditions. The reaction temperature may range from 125° C. to 350° C., e.g., from 200° C. to 325° C., from 225° C. to 300° C., or from 250° C. to 300° C. The pressure may range from 10 kPa to 3000 kPa, e.g., from 50 kPa to 2300 kPa, or from 100 kPa to 1500 kPa. The reactants may be fed to the reactor at a gas hourly space velocity (GHSV) of greater than 500 hr⁻¹, e.g., greater than 1000 hr⁻¹, greater than 2500 hr⁻¹ or even greater than 5000 hr⁻¹. In terms of ranges the GHSV may range from 50 hr⁻¹ to 50,000 hr⁻¹, e.g., from 500 hr⁻¹ to 30,000 hr⁻¹, from 1000 hr⁻¹ to 10,000 hr⁻¹, or from 1000 hr⁻¹ to 6500 hr⁻¹.

The hydrogenation optionally is carried out at a pressure just sufficient to overcome the pressure drop across the catalytic bed at the GHSV selected, although there is no bar to the use of higher pressures, it being understood that considerable pressure drop through the reactor bed may be experienced at high space velocities, e.g., 5000 hr⁻¹ or 6,500 hr⁻¹.

Although the reaction consumes two moles of hydrogen per mole of acetic acid to produce one mole of ethanol, the actual molar ratio of hydrogen to acetic acid in the feed stream may vary from about 100:1 to 1:100, e.g., from 50:1 to 1:50, from 20:1 to 1:2, or from 12:1 to 1:1. Most preferably, the molar ratio of hydrogen to acetic acid is greater than 2:1, e.g., greater than 4:1 or greater than 8:1.

Contact or residence time can also vary widely, depending upon such variables as amount of acetic acid, catalyst, reactor, temperature, and pressure. Typical contact times range from a fraction of a second to more than several hours when a catalyst system other than a fixed bed is used, with preferred contact times, at least for vapor phase reactions, of from 0.1 to 100 seconds, e.g., from 0.3 to 80 seconds or from 0.4 to 30 seconds.

The hydrogenation of acetic acid to form ethanol is preferably conducted in the presence of a hydrogenation catalyst. Suitable hydrogenation catalysts include catalysts comprising a first metal and optionally one or more of a second metal, a third metal or any number of additional metals, optionally on a catalyst support. The first and optional second and third metals may be selected from Group IB, IIB, IIIB, IVB, VB, VIIB, VIIB, VIII transition metals, a lanthanide metal, an actinide metal or a metal selected from any of Groups IIIA, IVA, VA, and VIA. Preferred metal combinations for some exemplary catalyst compositions include platinum/tin, platinum/ruthenium, platinum/rhenium, palladium/ruthenium, palladium/rhenium, cobalt/palladium, cobalt/platinum, cobalt/chromium, cobalt/ruthenium, cobalt/tin, silver/palladium, copper/palladium, copper/zinc, nickel/palladium, gold/palladium, ruthenium/rhenium, and ruthenium/iron. Exemplary catalysts are further described in U.S. Pat. No. 7,608,744 and U.S. Pub. No. 2010/0029995, the entireties of which are incorporated herein by reference. In another embodiment, the catalyst comprises a Co/Mo/S catalyst of the type described in U.S. Pub. No. 2009/0069609, the entirety of which is incorporated herein by reference.

In one embodiment, the catalyst comprises a first metal selected from the group consisting of copper, iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium, platinum, titanium, zinc, chromium, rhenium, molybdenum, and tungsten. Preferably, the first metal is selected from the group consisting of platinum, palladium, cobalt, nickel, and ruthenium. More preferably, the first metal is selected from platinum and palladium. In embodiments of the invention where the first metal comprises platinum, it is preferred that the catalyst comprises platinum in an amount less than 5 wt. %, e.g., less than 3 wt. % or less than 1 wt. %, due to the high commercial demand for platinum.

As indicated above, in some embodiments, the catalyst further comprises a second metal, which typically would function as a promoter. If present, the second metal preferably is selected from the group consisting of copper, molybdenum, tin, chromium, iron, cobalt, vanadium, tungsten, palladium, platinum, lanthanum, cerium, manganese, ruthenium, rhenium, gold, and nickel. More preferably, the second metal is selected from the group consisting of copper, tin, cobalt, rhenium, and nickel. More preferably, the second metal is selected from tin and rhenium.

In certain embodiments where the catalyst includes two or more metals, e.g., a first metal and a second metal, the first metal is present in the catalyst in an amount from 0.1 to 10 wt. %, e.g., from 0.1 to 5 wt. %, or from 0.1 to 3 wt. %. The second metal preferably is present in an amount from 0.1 to 20 wt. %, e.g., from 0.1 to 10 wt. %, or from 0.1 to 5 wt. %. For catalysts comprising two or more metals, the two or more metals may be alloyed with one another or may comprise a non-alloyed metal solution or mixture.

The preferred metal ratios may vary depending on the metals used in the catalyst. In some exemplary embodiments, the mole ratio of the first metal to the second metal is from 10:1 to 1:10, e.g., from 4:1 to 1:4, from 2:1 to 1:2, from 1.5:1 to 1:1.5 or from 1.1:1 to 1:1.1.

The catalyst may also comprise a third metal selected from any of the metals listed above in connection with the first or second metal, so long as the third metal is different from the first and second metals. In preferred aspects, the third metal is selected from the group consisting of cobalt, palladium, ruthenium, copper, zinc, platinum, tin, and rhenium. More preferably, the third metal is selected from cobalt, palladium, and ruthenium. When present, the total weight of the third metal preferably is from 0.05 to 4 wt. %, e.g., from 0.1 to 3 wt. %, or from 0.1 to 2 wt. %.

In addition to one or more metals, in some embodiments of the present invention the catalysts further comprise a support or a modified support. As used herein, the term “modified support” refers to a support that includes a support material and a support modifier, which adjusts the acidity of the support material.

The total weight of the support or modified support, based on the total weight of the catalyst, preferably is from 75 to 99.9 wt. %, e.g., from 78 to 97 wt. %, or from 80 to 95 wt. %. In preferred embodiments that utilize a modified support, the support modifier is present in an amount from 0.1 to 50 wt. %, e.g., from 0.2 to 25 wt. %, from 0.5 to 15 wt. %, or from 1 to 8 wt. %, based on the total weight of the catalyst. The metals of the catalysts may be dispersed throughout the support, layered throughout the support, coated on the outer surface of the support (i.e., egg shell), or decorated on the surface of the support.

As will be appreciated by those of ordinary skill in the art, support materials are selected such that the catalyst system is suitably active, selective and robust under the process conditions employed for the formation of ethanol.

Suitable support materials may include, for example, stable metal oxide-based supports or ceramic-based supports. Preferred supports include silicaceous supports, such as silica, silica/alumina, a Group IIA silicate such as calcium metasilicate, pyrogenic silica, high purity silica, and mixtures thereof. Other supports may include, but are not limited to, iron oxide, alumina, titania, zirconia, magnesium oxide, carbon, graphite, high surface area graphitized carbon, activated carbons, and mixtures thereof.

As indicated, the catalyst support may be modified with a support modifier. In some embodiments, the support modifier may be an acidic modifier that increases the acidity of the catalyst. Suitable acidic support modifiers may be selected from the group consisting of: oxides of Group IVB metals, oxides of Group VB metals, oxides of Group VIB metals, oxides of Group VIIB metals, oxides of Group VIIIB metals, aluminum oxides, and mixtures thereof. Acidic support modifiers include those selected from the group consisting of TiO₂, ZrO₂, Nb₂O₅, Ta₂O₅, Al₂O₃, B₂O₃, P₂O₅, and Sb₂O₃. Preferred acidic support modifiers include those selected from the group consisting of TiO₂, ZrO₂, Nb₂O₅, Ta₂O₅, and Al₂O₃. The acidic modifier may also include WO₃, MoO₃, Fe₂O₃, Cr₂O₃, V₂O₅, MnO₂, CuO, Co₂O₃, and Bi₂O₃.

In another embodiment, the support modifier may be a basic modifier that has a low volatility or no volatility. Such basic modifiers, for example, may be selected from the group consisting of: (i) alkaline earth oxides, (ii) alkali metal oxides, (iii) alkaline earth metal metasilicates, (iv) alkali metal metasilicates, (v) Group IIB metal oxides, (vi) Group IIB metal metasilicates, (vii) Group IIIB metal oxides, (viii) Group IIIB metal metasilicates, and mixtures thereof. In addition to oxides and metasilicates, other types of modifiers including nitrates, nitrites, acetates, and lactates may be used. Preferably, the support modifier is selected from the group consisting of oxides and metasilicates of any of sodium, potassium, magnesium, calcium, scandium, yttrium, and zinc, as well as mixtures of any of the foregoing. More preferably, the basic support modifier is a calcium silicate, and even more preferably calcium metasilicate (CaSiO₃). If the basic support modifier comprises calcium metasilicate, it is preferred that at least a portion of the calcium metasilicate is in crystalline form.

A preferred silica support material is SS61138 High Surface Area (HSA) Silica Catalyst Carrier from Saint Gobain N or Pro. The Saint-Gobain N or Pro SS61138 silica exhibits the following properties: contains approximately 95 wt. % high surface area silica; surface area of about 250 m²/g; median pore diameter of about 12 nm; average pore volume of about 1.0 cm³/g as measured by mercury intrusion porosimetry and a packing density of about 0.352 g/cm³ (22 lb/ft³).

A preferred silica/alumina support material is KA-160 silica spheres from Sud Chemie having a nominal diameter of about 5 mm, a density of about 0.562 g/ml, an absorptivity of about 0.583 g H₂O/g support, a surface area of about 160 to 175 m²/g, and a pore volume of about 0.68 ml/g.

The catalyst compositions suitable for use with the present invention preferably are formed through metal impregnation of the modified support, although other processes such as chemical vapor deposition may also be employed. Such impregnation techniques are described in U.S. Pat. Nos. 7,608,744 and 7,863,489 and U.S. Pub. No. 2010/0197485 referred to above, the entireties of which are incorporated herein by reference.

In particular, the hydrogenation of acetic acid may achieve favorable conversion of acetic acid and favorable selectivity and productivity to ethanol. For purposes of the present invention, the term “conversion” refers to the amount of acetic acid in the feed that is converted to a compound other than acetic acid. Conversion is expressed as a mole percentage based on acetic acid in the feed. The conversion may be at least 10%, e.g., at least 20%, at least 40%, at least 50%, at least 60%, at least 70% or at least 80%. Although catalysts that have high conversions are desirable, such as at least 80% or at least 90%, in some embodiments a low conversion may be acceptable at high selectivity for ethanol. It is, of course, well understood that in many cases, it is possible to compensate for conversion by appropriate recycle streams or use of larger reactors, but it is more difficult to compensate for poor selectivity.

Selectivity is expressed as a mole percent based on converted acetic acid. It should be understood that each compound converted from acetic acid has an independent selectivity and that selectivity is independent from conversion. For example, if 60 mole % of the converted acetic acid is converted to ethanol, we refer to the ethanol selectivity as 60%. Preferably, the catalyst selectivity to ethoxylates is at least 60%, e.g., at least 70%, or at least 80%. As used herein, the term “ethoxylates” refers specifically to the compounds ethanol, acetaldehyde, and ethyl acetate. Preferably, the selectivity to ethanol is at least 80%, e.g., at least 85% or at least 88%. Preferred embodiments of the hydrogenation process also have low selectivity to undesirable products, such as methane, ethane, and carbon dioxide. The selectivity to these undesirable products preferably is less than 4%, e.g., less than 2% or less than 1%. More preferably, these undesirable products are present in undetectable amounts. Formation of alkanes may be low, and ideally less than 2%, less than 1%, or less than 0.5% of the acetic acid passed over the catalyst is converted to alkanes, which have little value other than as fuel.

The term “productivity,” as used herein, refers to the grams of a specified product, e.g., ethanol, formed during the hydrogenation based on the kilograms of catalyst used per hour. A productivity of at least 100 grams of ethanol per kilogram of catalyst per hour, e.g., at least 400 grams of ethanol per kilogram of catalyst per hour or at least 600 grams of ethanol per kilogram of catalyst per hour, is preferred. In terms of ranges, the productivity preferably is from 100 to 3,000 grams of ethanol per kilogram of catalyst per hour, e.g., from 400 to 2,500 grams of ethanol per kilogram of catalyst per hour or from 600 to 2,000 grams of ethanol per kilogram of catalyst per hour.

Operating under the conditions of the present invention may result in ethanol production on the order of at least 0.1 tons of ethanol per hour, e.g., at least 1 ton of ethanol per hour, at least 5 tons of ethanol per hour, or at least 10 tons of ethanol per hour. Larger scale industrial production of ethanol, depending on the scale, generally should be at least 1 ton of ethanol per hour, e.g., at least 15 tons of ethanol per hour or at least 30 tons of ethanol per hour. In terms of ranges, for large scale industrial production of ethanol, the process of the present invention may produce from 0.1 to 160 tons of ethanol per hour, e.g., from 15 to 160 tons of ethanol per hour or from 30 to 80 tons of ethanol per hour. Ethanol production from fermentation, due the economies of scale, typically does not permit the single facility ethanol production that may be achievable by employing embodiments of the present invention.

In various embodiments of the present invention, the crude ethanol product produced by the hydrogenation process, before any subsequent processing, such as purification and separation, will typically comprise unreacted acetic acid, ethanol and water. As used herein, the term “crude ethanol product” refers to any composition comprising from 5 to 70 wt. % ethanol and from 5 to 40 wt. % water. Exemplary component ranges for the crude ethanol product are provided in Table 1. The others identified in Table 1 may include, for example, esters, ethers, aldehydes, ketones, alkanes, and carbon dioxide.

TABLE 1
CRUDE ETHANOL PRODUCT COMPOSITIONS
	Conc.	Conc.	Conc.	Conc.
Component	(wt. %)	(wt. %)	(wt. %)	(wt. %)
Ethanol	5 to 70	15 to 70	15 to 50	25 to 50
Acetic Acid	0 to 90	0 to 50	15 to 70	20 to 70
Water	5 to 40	5 to 30	10 to 30	10 to 26
Ethyl Acetate	0 to 30	0 to 20	1 to 12	3 to 10
Acetaldehyde	0 to 10	0 to 3	0.1 to 3	0.2 to 2
Others	0.1 to 10	0.1 to 6	0.1 to 4	—

In one embodiment, the crude ethanol product may comprise acetic acid in an amount less than 20 wt. %, e.g., less than 15 wt. %, less than 10 wt. % or less than 5 wt. %. In embodiments having lower amounts of acetic acid, the conversion of acetic acid is preferably greater than 75%, e.g., greater than 85% or greater than 90%. In addition, the selectivity to ethanol may also be preferably high, and is preferably greater than 75%, e.g., greater than 85% or greater than 90%.

Ethanol Recovery

Exemplary ethanol recovery systems in accordance with embodiments of the present invention are shown in FIGS. 1 and 2. System 100 comprises reaction zone 101 and separation zone 102. Hydrogen and acetic acid are fed to a vaporizer 110 via lines 104 and 105, respectively, to create a vapor feed stream in line 111 that is directed to reactor 103. In one embodiment, lines 104 and 105 may be combined and jointly fed to the vaporizer 110. The temperature of the vapor feed stream in line 111 is preferably from 100° C. to 350° C., e.g., from 120° C. to 310° C. or from 150° C. to 300° C. Any feed that is not vaporized is removed from vaporizer 110, as shown in FIG. 1, and may be recycled or discarded. In addition, although FIG. 1 shows line 111 being directed to the top of reactor 103, line 111 may be directed to the side, upper portion, or bottom of reactor 103. Further modifications and additional components to reaction zone 101 and separation zone 102 are described below.

Reactor 103 contains the catalyst that is used in the hydrogenation of the carboxylic acid, preferably acetic acid. In one embodiment, one or more guard beds (not shown) may be used upstream of the reactor, optionally upstream of vaporizer 110, to protect the catalyst from poisons or undesirable impurities contained in the feed or return/recycle streams. Such guard beds may be employed in the vapor or liquid streams. Suitable guard bed materials may include, for example, carbon, silica, alumina, ceramic, or resins. In one aspect, the guard bed media is functionalized, e.g., silver functionalized, to trap particular species such as sulfur or halogens. During the hydrogenation process, a crude ethanol product stream is withdrawn, preferably continuously, from reactor 103 via line 112.

The crude ethanol product stream in line 112 may be condensed and fed to a separator 106, which, in turn, provides a vapor stream 113 and a liquid stream 114. Suitable separators 106 include a flasher or a knockout pot. The separator 106 may operate at a temperature of from 20° C. to 250° C., e.g., from 30° C. to 225° C. or from 60° C. to 200° C. The pressure of separator 106 may be from 50 kPa to 2000 kPa, e.g., from 75 kPa to 1500 kPa or from 100 kPa to 1000 kPa. Optionally, the crude ethanol product in line 112 may pass through one or more membranes to separate hydrogen and/or other non-condensable gases.

The vapor stream 113 exiting separator 106 may comprise hydrogen and hydrocarbons, which may be purged and/or returned to reaction zone 101. As shown, vapor stream 113 is combined with the hydrogen feed 104 and co-fed to vaporizer 110. In some embodiments, the returned vapor stream 113 may be compressed before being combined with hydrogen feed 104.

The liquid stream 114 from separator 106 is withdrawn and pumped to the side of distillation column 107. In one embodiment, the contents of liquid stream 114 are substantially similar to the crude ethanol product obtained from the reactor in line 112, except that the composition has been depleted of hydrogen, carbon dioxide, methane and/or ethane, which are preferably removed by separator 106. Accordingly, liquid stream 114 may also be referred to as a crude ethanol product. Exemplary components of liquid stream 114 are provided in Table 2. It should be understood that liquid stream 114 may contain other components, not listed, such as components derived from the feed.

TABLE 2
COMPOSITION OF LIQUID STREAM 114
	Conc. (wt. %)	Conc. (wt. %)	Conc. (wt. %)
Ethanol	5 to 70	10 to 60	15 to 50
Acetic Acid	<90	5 to 80	15 to 70
Water	5 to 40	5 to 30	10 to 30
Ethyl Acetate	<30	0.001 to 20	1 to 12
Acetaldehyde	<10	0.001 to 3	0.1 to 3
Acetal	<5	0.001 to 2	0.005 to 1
Acetone	<5	0.0005 to 0.05	0.001 to 0.03
Other Esters	<5	<0.005	<0.001
Other Ethers	<5	<0.005	<0.001
Other Alcohols	<5	<0.005	<0.001

The amounts indicated as less than (<) in the tables throughout present application are preferably not present and if present may be present in trace amounts or in amounts greater than 0.0001 wt. %.

The “other esters” in Table 2 may include, but are not limited to, ethyl propionate, methyl acetate, isopropyl acetate, n-propyl acetate, n-butyl acetate or mixtures thereof. The “other ethers” in Table 2 may include, but are not limited to, diethyl ether, methyl ethyl ether, isobutyl ethyl ether or mixtures thereof. The “other alcohols” in Table 2 may include, but are not limited to, methanol, isopropanol, n-propanol, n-butanol or mixtures thereof. In one embodiment, the liquid stream 114, may comprise propanol, e.g., isopropanol and/or n-propanol, in an amount from 0.001 to 0.1 wt. %, from 0.001 to 0.05 wt. % or from 0.001 to 0.03 wt. %. It should be understood that these other components may be carried through in any of the distillate or residue streams described herein and will not be further described herein, unless indicated otherwise.

Optionally, crude ethanol product in line 112 or in liquid stream 114 may be further fed to an esterification reactor, hydrogenolysis reactor, or combination thereof. An esterification reactor may be used to consume acetic acid present in the crude ethanol product to further reduce the amount of acetic acid to be removed. Hydrogenolysis may be used to convert ethyl acetate in the crude ethanol product to ethanol.

In the embodiment shown in FIG. 1, line 114 is introduced in the middle part of first column 107, e.g., second or third quarter. Depending on the composition of the crude ethanol product in line 114 and the operating conditions of first column 107, the first column 107 separates the crude ethanol product in line 114, preferably continuously, into a first distillate in line 117 and a first residue in line 116. In one embodiment, no entrainers are added to first column 107. The first distillate in line 117 comprises ethanol, other organics and water. The first residue in line 116 comprises unreacted acetic acid, water, and other heavy components, if present. In some embodiments, especially with higher conversions of acetic acid of at least 80%, or at least 90%, it may be beneficial to remove a substantial portion of the water, e.g. at least 35% to 90% of the water, from liquid stream 114, in the first residue along with substantially all the acetic acid.

When column 107 is operated under about 170 kPa, the temperature of the residue exiting in line 116 preferably is from 90° C. to 130° C., e.g., from 95° C. to 120° C. or from 100° C. to 115° C. The temperature of the distillate exiting in line 117 preferably is from 60° C. to 90° C., e.g., from 65° C. to 85° C. or from 70° C. to 80° C. In some embodiments, the pressure of first column 107 may range from 0.1 kPa to 510 kPa, e.g., from 1 kPa to 475 kPa or from 1 kPa to 375 kPa.

The first distillate in line 117 comprises water, as discussed above, in addition to ethanol and other organics. In terms of ranges, the concentration of water in the first distillate in line 117 preferably is from 4 wt. % to 38 wt. %, e.g., from 7 wt. % to 32 wt. %, or from 7 to 25 wt. %. As shown in FIG. 1, prior to condensing the first distillate in line 117, first distillate in line 117 is fed to a water separator 118 to dehydrate the first distillate. Water separator 118 may be an adsorption unit, membrane, molecular sieves, extractive column distillation, or a combination thereof.

In a preferred embodiment, water separator 118 is a pressure swing adsorption (PSA) unit. The PSA unit is optionally operated at a temperature from 30° C. to 160° C., e.g., from 80° C. to 140° C., and a pressure of from 0.01 kPa to 550 kPa, e.g., from 1 kPa to 150 kPa. The PSA unit may comprise two to five beds. Water separator 118 may remove at least 95% of the water from the first distillate in line 117, and more preferably from 99% to 99.99% of the water from the first distillate, in a water stream 119. All or a portion of water stream 119 may be returned to column 107, where it preferably is ultimately recovered from column 107 in the first residue in line 116. Additionally or alternatively, all or a portion of water stream 119 may be purged via line 115. The remaining portion of first distillate 117 exits the water separator 118 as ethanol mixture stream 120. A portion of the first distillate in line 108 may be condensed and refluxed to first column 107, as shown, for example, at a ratio of from 10:1 to 1:100, e.g., from 2:1 to 1:50 or from 1:1 to 1:10. It is understand that reflux ratios may vary with the number of stages, feed locations, column efficiency and/or feed composition. Operating with a reflux ratio of greater than 3:1 may be less preferred because more energy may required to operate the first column 107. Preferably, ethanol mixture stream 120 is not returned or refluxed to first column 107.

In another aspect, shown in FIG. 1, all or a portion of the water in water stream 119 is fed to first column 107 at a point below where the first distillate in line 108 is refluxed to first column 107.

Exemplary components of ethanol mixture stream 120 and first residue in line 116 are provided in Table 3 below. It should also be understood that these streams may also contain other components, not listed, such as components derived from the feed.

TABLE 3
FIRST COLUMN WITH PSA
	Conc.
	(wt. %)	Conc. (wt. %)	Conc. (wt. %)
Ethanol Mixture Stream
Ethanol	20 to 95	30 to 95	40 to 95
Water	<10	0.01 to 6	0.1 to 2
Acetic Acid	<2	0.001 to 0.5	0.01 to 0.2
Ethyl Acetate	<60	1 to 55	5 to 55
Acetaldehyde	<10	0.001 to 5	0.01 to 4
Acetal	<0.1	<0.1	<0.05
Acetone	<0.05	0.001 to 0.03	0.01 to 0.025
Residue
Acetic Acid	<90	1 to 50	2 to 35
Water	30 to 100	45 to 95	60 to 90
Ethanol	<1	<0.9	<0.3

Some species, such as acetals, may decompose in first column 107 such that very low amounts, or even no detectable amounts, of acetals remain in the distillate or residue. In addition, an equilibrium reaction between acetic acid and ethanol or between ethyl acetate and water may occur in the crude ethanol product after it exits reactor 103. Depending on the concentration of acetic acid in the crude ethanol product, this equilibrium may be driven toward formation of ethyl acetate. This equilibrium may be regulated using the residence time and/or temperature of crude ethanol product.

Depending on the amount of water and acetic acid contained in the residue of first column 107, line 116 may be treated in one or more of the following processes. The following are exemplary processes for further treating first residue and it should be understood that any of the following may be used regardless of acetic acid concentration. When the residue comprises a majority of acetic acid, e.g., greater than 70 wt. %, the residue may be recycled to the reactor without any separation of the water. In one embodiment, the residue may be separated into an acetic acid stream and a water stream when the residue comprises a majority of acetic acid, e.g., greater than 50 wt. %. Acetic acid may also be recovered in some embodiments from first residue having a lower acetic acid concentration. The residue may be separated into the acetic acid and water streams by a distillation column or one or more membranes. If a membrane or an array of membranes is employed to separate the acetic acid from the water, the membrane or array of membranes may be selected from any suitable acid resistant membrane that is capable of removing a permeate water stream. The resulting acetic acid stream optionally is returned to reactor 103. The resulting water stream may be used as an extractive agent, for example to second column 109 via line 121 as discussed below, or to hydrolyze an ester-containing stream in a hydrolysis unit.

In other embodiments, for example where first residue in line 116 comprises less than 50 wt. % acetic acid, possible options include one or more of: (i) returning a portion of the residue to reactor 103, (ii) neutralizing the acetic acid, (iii) reacting the acetic acid with an alcohol, or (iv) disposing of the residue in a waste water treatment facility. It also may be possible to separate a residue comprising less than 50 wt. % acetic acid using a weak acid recovery (WAR) distillation column to which a solvent (optionally acting as an azeotroping agent) may be added. Exemplary solvents that may be suitable for this purpose include ethyl acetate, propyl acetate, isopropyl acetate, butyl acetate, vinyl acetate, diisopropyl ether, carbon disulfide, tetrahydrofuran, isopropanol, ethanol, and C₃-C₁₂ alkanes. When neutralizing the acetic acid, it is preferred that the residue in line 116 comprises less than 10 wt. % acetic acid. Acetic acid may be neutralized with any suitable alkali or alkaline earth metal base, such as sodium hydroxide or potassium hydroxide. When reacting acetic acid with an alcohol, it is preferred that the residue comprises less than 50 wt. % acetic acid. The alcohol may be any suitable alcohol, such as methanol, ethanol, propanol, butanol, or mixtures thereof. The reaction forms an ester that may be integrated with other systems, such as carbonylation production or an ester production process. Preferably, the alcohol comprises ethanol and the resulting ester comprises ethyl acetate. Optionally, the resulting ester may be fed to the hydrogenation reactor.

In some embodiments, when the first residue comprises very minor amounts of acetic acid, e.g., less than 5 wt. %, the residue may be disposed of to a waste water treatment facility without further processing. The organic content, e.g., acetic acid content, of the residue beneficially may be suitable to feed microorganisms used in the waste water treatment facility.

Reverting to FIG. 1, ethanol mixture stream 120 is introduced to a second column 109, also referred to as the “light ends column,” preferably in the top part of column 109, e.g., top half or top third. Second column 109 may be a tray column or packed column. In one embodiment, second column 109 is a tray column having from 5 to 70 trays, e.g., from 15 to 50 trays or from 20 to 45 trays. As one example, when a 30 tray column is utilized in a column without water extraction, ethanol mixture stream 120 is introduced at tray 2.

Optionally, the light ends column may be an extractive distillation column Suitable extractive agents may include, for example, dimethylsulfoxide, glycerine, diethylene glycol, 1-naphthol, hydroquinone, N,N′-dimethylformamide, 1,4-butanediol; ethylene glycol-1,5-pentanediol; propylene glycol-tetraethylene glycol-polyethylene glycol; glycerine-propylene glycol-tetraethylene glycol-1,4-butanediol, ethyl ether, methyl formate, cyclohexane, N,N-dimethyl-1,3-propanediamine, N,N′-dimethylethylenediamine, diethylene triamine, hexamethylene diamine and 1,3-diaminopentane, an alkylated thiopene, dodecane, tridecane, tetradecane, chlorinated paraffins, or a combination thereof. In another aspect, the extractive agent may comprise water. If the extraction agent comprises water, the water may be obtained from an external source or from an internal return/recycle line from one or more of the other columns, such as the water stream 119. Generally, the extractive agent is fed above the entry point of ethanol mixture stream 120, as shown by optional line 121. When extractive agents are used, a suitable recovery system, such as a further distillation column, may be used to remove the extractive agent and recycle the extractive agent if desired.

Second column 109 is operated to separate ethanol mixture stream 120 or a portion thereof into a second distillate in line 123 and a second residue in line 122. Second distillate in line 123 may comprise, for example, ethyl acetate and acetaldehyde, while second residue comprises ethanol, which optionally is a salable finished ethanol product.

Although the temperature and pressure of second column 109 may vary, when at about 20 kPa to 70 kPa, the temperature of the second residue exiting in line 122 preferably is from 30° C. to 75° C., e.g., from 35° C. to 70° C. or from 40° C. to 65° C. The temperature of the second distillate exiting in line 123 preferably is from 20° C. to 55° C., e.g., from 25° C. to 50° C. or from 30° C. to 45° C. Second column 109 may operate at a reduced pressure, near or at vacuum conditions, to further favor separation of ethyl acetate and ethanol. In other embodiments, the pressure of second column 109 may range from 0.1 kPa to 510 kPa, e.g., from 1 kPa to 475 kPa or from 1 kPa to 375 kPa. Exemplary distillate and residue compositions for second column 109 are provided in Table 4, below. It should be understood that the second distillate and second residue may also contain other components, not listed, such as components in the feed.

TABLE 4
SECOND COLUMN
	Conc. (wt. %)	Conc. (wt. %)	Conc. (wt. %)
Second Distillate
Ethyl Acetate	5 to 90	10 to 80	15 to 75
Acetaldehyde	<60	1 to 40	1 to 35
Ethanol	<45	0.001 to 40	0.01 to 35
Water	<20	0.01 to 10	0.1 to 5
Second Residue
Ethanol	80 to 99.5	85 to 97	60 to 95
Water	<20	0.001 to 15	0.01 to 10
Ethyl Acetate	<1	0.001 to 2	0.001 to 0.5
Acetic Acid	<0.5	<0.01	0.001 to 0.01

The weight ratio of ethanol in the second residue to ethanol in the second distillate preferably is at least 2:1, e.g., at least 5:1, at least 8:1, at least 10:1 or at least 15:1. The weight ratio of ethyl acetate in the second residue to ethyl acetate in the second distillate preferably is less than 0.4:1, e.g., less than 0.2:1 or less than 0.1:1.

Second distillate in line 123, which comprises ethyl acetate and/or acetaldehyde, preferably is refluxed as shown in FIG. 1, for example, at a reflux ratio of from 1:30 to 30:1, e.g., from 1:15 to 15:1 or from 1:5 to 5:1. In one aspect, the second distillate in line 123 or a portion thereof may be returned reactor 103. For example, it may be advantageous to return a portion of second distillate 123 to reactor 103. The ethyl acetate and/or acetaldehyde in the second distillate may be further reacted in hydrogenation reactor 103 or in an secondary reactor. The outflow from the secondary reactor may be fed to reactor 103 to produce additional ethanol or to a distillation column, such as columns, 107, 115, or 118, to recover additional ethanol.

In the embodiment shown in FIG. 1, the finished ethanol product produced by the process of the present invention may be taken from the second residue in line 122. Advantageously, the finished ethanol product may be recovered using water separator and two columns according to the present invention. The finished ethanol product may be an industrial grade ethanol comprising from 75 to 96 wt. % ethanol, e.g., from 80 to 96 wt. % or from 85 to 96 wt. % ethanol, based on the total weight of the ethanol product. Exemplary finished ethanol compositional ranges are provided below in Table 5.

TABLE 5
FINISHED ETHANOL COMPOSITIONS
Component	Conc. (wt. %)	Conc. (wt. %)	Conc. (wt. %)
Ethanol	75 to 96	80 to 96	85 to 96
Water	<12	1 to 9	3 to 8
Acetic Acid	<1	<0.1	<0.01
Ethyl Acetate	<2	<0.5	<0.05
Acetal	<0.05	<0.01	<0.005
Acetone	<0.05	<0.01	<0.005
Isopropanol	<0.5	<0.1	<0.05
n-propanol	<0.5	<0.1	<0.05

The finished ethanol composition of the present invention preferably contains very low amounts, e.g., less than 0.5 wt. %, of other alcohols, such as methanol, butanol, isobutanol, isoamyl alcohol and other C₄-C₂₀ alcohols. In one embodiment, the amount of isopropanol in the finished ethanol composition is from 80 to 1,000 wppm, e.g., from 95 to 1,000 wppm, from 100 to 700 wppm, or from 150 to 500 wppm. In one embodiment, the finished ethanol composition is substantially free of acetaldehyde, optionally comprising less than 8 wppm acetaldehyde, e.g., less than 5 wppm or less than 1 wppm.

In some embodiments, when further water separation is used, the ethanol product may be withdrawn as a stream from the water separation unit as discussed above. In such embodiments, the ethanol concentration of the ethanol product may be higher than indicated in Table 5, and preferably is greater than 97 wt. % ethanol, e.g., greater than 98 wt. % or greater than 99.5 wt. %. The ethanol product in this aspect preferably comprises less than 3 wt. % water, e.g., less than 2 wt. % or less than 0.5 wt. %.

In another embodiment of the present invention, as shown in FIG. 2, the first distillate in line 117 from first column 107 passes through a compressor 130 and is fed to a membrane 131. Membrane 131 preferably operates in the vapor phase. The water in the first distillate in line 117 permeates across the membrane 131 to form a permeate stream 132. Permeate stream 132 may comprise at least 80% of the water from the first distillate, and more preferably at least 90% of the water. In one embodiment, permeate stream 132 also comprises less than 10% of the ethanol from the first distillate, e.g., less than 5% of the ethanol. A portion of permeate stream 132 may be returned to column 107, preferably below the feed point of the liquid stream 114. A portion of permeate stream 132 may be purged from the system, as shown by line 133, to control the amount of water in column 107.

The portion of distillate that does not cross membrane 131 forms retentate stream 134, e.g., an ethanol mixture stream, and is fed to second column 109 as described above in FIG. 1. In one embodiment, retentate stream 134 comprises more water than shown in the ethanol mixture stream compositions provided in Table 3. For example, retentate stream 134 may comprise from 0 to 10 wt. % water, and more preferably form 0 to 5 wt. % water. The additional water present in retentate stream 134 may be beneficial in separating ethyl acetate and ethanol in second column 109. The additional water that is fed to second column 109 preferably is removed with the second residue in line 122 and may necessitate a secondary dewatering step, as discussed above, to provide an ethanol product having the desired water concentration.

Although one membrane is shown in FIG. 2, it should be understood that a suitable array of membranes may also be used.

In most embodiments of the present invention, it is desirable to remove water from the first distillate as shown in FIGS. 1 and 2. In FIG. 3, the first distillate in line 117 is split into a main line 140 and a bypass line 141. In one embodiment, the split ratio may range from 10:1 to 1:10, e.g., from 5:1 to 1:5 or about 1:1. Main line 140 is fed to water separator 118 to produce a water stream 142 and an ethanol mixture stream 143. Water stream 142 may be returned to first column 107 as described above in FIGS. 1 and 2. A portion of first distillate in line 117 may be condensed and refluxed to first column 107 via line 108, e.g., at a reflux ratio from 10:1 to 1:100, e.g., from 2:1 to 1:50 or from 1:1 to 1:10. A bypass line 141 may be taken from the reflux line 108 and fed directly to second column 109 along with ethanol mixture stream 143. Flow control valves (not shown) may be used to control the split between the main line 140 and bypass line 141 such that a desired water concentration is fed to the second column 109. In one embodiment, not shown, a portion of bypass line 141 is refluxed to first column 107. In another embodiment, bypass line 141 is not refluxed to first column 107 and there may be a separate reflux line (not shown) A portion of bypass line 141 may be condensed and separately fed to second column 109 with all or a portion of ethanol mixture stream 143.

The separation of the first distillate 117 into main line 140 and bypass line 141 may be controlled based on the desired concentration of water of the combined streams 140 and 141. In an embodiment, bypass line 141 contains a higher concentration of water than the ethanol mixture stream 143. The combined distillate and ethanol mixture has a total water concentration of greater than 0.5 wt. %, e.g., greater than 2 wt. % or greater than 5 wt. %. In terms of ranges, the total water concentration of the combined distillate and ethanol mixture may be from 0.5 to 15 wt. %, e.g., from 2 to 12 wt. %, or from 5 to 10 wt. %. The water content in bypass line 141 may be beneficial for the separation of ethanol and ethyl acetate in distillation column 109. In another embodiment, additional water may be added to column 109 via line 121. Optionally, portions of the water in water stream 142 may be added to column 109.

Each of the columns shown in FIGS. 1-3 may be selected from any distillation column capable of performing the specified separation and/or purification step. Each column preferably comprises a tray column having from 1 to 150 trays, e.g., from 10 to 100 trays, from 20 to 95 trays or from 30 to 75 trays. The trays may be sieve trays, fixed valve trays, movable valve trays, or any other suitable design known in the art. In other embodiments, a packed column may be used. For packed columns, structured packing or random packing may be employed. The trays or packing may be arranged in one continuous column or they may be arranged in two or more columns such that the vapor from the first section enters the second section while the liquid from the second section enters the first section, etc.

For convenience, the distillate and residue of the first column may also be referred to as the “first distillate” or “first residue.” The distillates or residues of the other columns may also be referred to with similar numeric modifiers (second, third, etc.) in order to distinguish them from one another, but such modifiers should not be construed as requiring any particular separation order.

The associated condensers and liquid separation vessels that may be employed with each of the distillation columns may be of any conventional design and are simplified in the figures. Heat may be supplied to the base of each column or to a circulating bottoms stream through a heat exchanger or reboiler. Other types of reboilers, such as internal reboilers, may also be used. The heat that is provided to the reboilers may be derived from any heat generated during the process that is integrated with the reboilers or from an external source such as another heat generating chemical process or a boiler. Although one reactor and one flasher are shown in the figures, additional reactors, flashers, condensers, heating elements, and other components may be used in various embodiments of the present invention. As will be recognized by those skilled in the art, various condensers, pumps, compressors, reboilers, drums, valves, connectors, separation vessels, etc., normally employed in carrying out chemical processes may also be combined and employed in the processes of the present invention.

The temperatures and pressures employed in the columns may vary. As a practical matter, pressures from 10 kPa to 3000 kPa will generally be employed in these zones although in some embodiments subatmospheric pressures or superatmospheric pressures may be employed. Temperatures within the various zones will normally range between the boiling points of the composition removed as the distillate and the composition removed as the residue. As will be recognized by those skilled in the art, the temperature at a given location in an operating distillation column is dependent on the composition of the material at that location and the pressure of column. In addition, feed rates may vary depending on the size of the production process and, if described, may be generically referred to in terms of feed weight ratios.

The finished ethanol composition produced by the embodiments of the present invention may be used in a variety of applications including applications as fuels, solvents, chemical feedstocks, pharmaceutical products, cleansers, sanitizers, hydrogenation transport or consumption. In fuel applications, the finished ethanol composition may be blended with gasoline for motor vehicles such as automobiles, boats and small piston engine aircraft. In non-fuel applications, the finished ethanol composition may be used as a solvent for toiletry and cosmetic preparations, detergents, disinfectants, coatings, inks, and pharmaceuticals. The finished ethanol composition may also be used as a processing solvent in manufacturing processes for medicinal products, food preparations, dyes, photochemicals and latex processing.

The finished ethanol composition may also be used as a chemical feedstock to make other chemicals such as vinegar, ethyl acrylate, ethyl acetate, ethylene, glycol ethers, ethylamines, aldehydes, and higher alcohols, especially butanol. In the production of ethyl acetate, the finished ethanol composition may be esterified with acetic acid. In another application, the finished ethanol composition may be dehydrated to produce ethylene. Any known dehydration catalyst can be employed to dehydrate ethanol, such as those described in copending U.S. Pub. Nos. 2010/0030002 and 2010/0030001, the entire contents and disclosures of which are hereby incorporated by reference. A zeolite catalyst, for example, may be employed as the dehydration catalyst. Preferably, the zeolite has a pore diameter of at least about 0.6 nm, and preferred zeolites include dehydration catalysts selected from the group consisting of mordenites, ZSM-5, a zeolite X and a zeolite Y. Zeolite X is described, for example, in U.S. Pat. No. 2,882,244 and zeolite Y in U.S. Pat. No. 3,130,007, the entireties of which are hereby incorporated herein by reference.

In order that the invention disclosed herein may be more efficiently understood, an example is provided below. It should be understood that these examples are for illustrative purposes only and is not to be construed as limiting the invention in any manner.

EXAMPLES

The following examples were prepared with ASPEN Plus 7.1 simulation software to test various feed composition and separation systems.

Example 1

Reactor product comprising 56 wt. % ethanol, 38 wt. % water, 2 wt. % ethyl acetate, 2 wt. % acetaldehyde, 1 wt. % acetic acid, and 1 wt. % other organics were fed at 10 lb per hour into a distillation column. The distillation column operated with 50 trays at atmospheric pressure. The feed tray was located at 35^th tray from the top. Table 6 compares the distillate composition from two different runs. In Run A the energy requirements per ton of ethanol refined was about 5 times greater than Run B.

	TABLE 6
	Distillate	Run A	Run B
	Ethanol	86.6 wt. %	75.5 wt. %
	Water	7.2 wt. %	19.1 wt. %
	Ethyl acetate	3.1 wt. %	2.7 wt. %
	Acetaldehyde	3.1 wt. %	2.7 wt. %

Example 2

The distillate obtained from Run B was further processed using a water separator. As shown in Table 7, both the 3A zeolite molecular sieve and zeolite membrane achieved significant removal of water to produce an ethanol mixture that can be further separated to recover ethanol.

	TABLE 7
	Water Separator	Molecular Sieve	Membrane
	Ethanol Mixture
	Ethanol	93.2 wt. %	89.9 wt. %
	Water	0.2 wt. %	3.8 wt. %
	Ethyl Acetate	3.3 wt. %	3.2 wt. %
	Acetaldehyde	3.3 wt. %	3.2 wt. %
	% Water Removed	99%	97%

Example 3

Reactor product comprising 56 mol. % ethanol, 38 mol. % water, 2 mol. % ethyl acetate, 2 mol. % acetaldehyde, 1 mol. % acetic acid, and 1 mol. % other organics were fed to a distillation column having an PSA to process a split of the distillate. The distillation column operated with 37 trays and the feed tray was located at 17^th tray from the top. The distillation operated 135.8 kPa with a temperature in the overhead of 89° C. and in the residue in 120° C. The distillate was split into equal portions (1:1) and one portion of distillate was fed to a PSA unit operated at a temperature of 52° C. a pressure of 34.5 kPa. The other portion of the distillate was combined with the dry ethanol stream to form the second feed that may be processed in a light-ends column.

As shown in Table 8, the energy requirements for the column varied with the reflux ratios. Energy is based on the MMBtu per ton of ethanol refined in the column Run C, with the highest reflux ratio, demonstrated the highest energy. More water was carried through to the second feed at lower reflux ratios, e.g., Runs D, E, and F, however, the energy savings in this column is sufficient to allow further separation. In addition, the presence of water in the light-ends column may be beneficial in separating the light organics, ethyl acetate and acetaldehyde, from ethanol. Run G and H demonstrated low energy based on the simulation from ASPEN, but the amount of acetic acid was also greater than Runs D, E, and F.

TABLE 8
High Conversion of Acetic Acid (99%)
	Run C	Run D	Run E	Run F	Run G	Run H
Distillate - wt. %
Ethanol	84.9	83.6	80.8	74.3	73.6	73.1
Water	7.4	8.9	12.0	19.0	19.8	20.4
Ethyl Acetate	3.0	3.0	2.9	2.7	2.6	2.6
Acetaldehyde	2.3	2.2	2.2	2.0	2.0	2.0
Acetic Acid	<0.01	<0.01	<0.01	<0.01	0.03	0.06
Split	1:1	1:1	1:1	1:1	1:1	1:1
Reflux Ratio	5:1	2:1	3:4	1:10	1:20	1:100
Energy	9.74	5.06	3.18	2.33	2.26	2.20
(MMBtu/ton
ETOH)
Dry Ethanol Stream - wt. %
Ethanol	91.7	91.7	91.7	91.6	91.5	91.5
Water	0.1	0.1	0.1	0.2	0.2	0.3
Ethyl Acetate	3.3	3.3	3.3	3.3	3.3	3.3
Acetaldehyde	2.5	2.5	2.5	2.5	2.5	2.5
Water Stream - wt. %
Water	98.8	99.0	99.3	99.6	99.6	99.6
Ethanol	1.1	0.9	0.7	0.4	0.4	0.4
Second Feed - wt. %
Ethanol	89.7	88.9	87.3	83.4	83.0	82.6
Water	3.6	4.4	6.2	10.3	10.8	11.2
Ethyl Acetate	3.2	3.1	3.1	2.9	2.9	2.9
Acetaldehyde	3.1	3.0	3.0	2.9	2.8	2.8

The amount of water fed to the second column may be controlled by varying the split of the distillate that is fed to the PSA, as opposing to increasing the energy needed in the column to remove water from the distillate.

While the invention has been described in detail, modifications within the spirit and scope of the invention will be readily apparent to those of skill in the art. In addition, it should be understood that aspects of the invention and portions of various embodiments and various features recited herein and/or in the appended claims may be combined or interchanged either in whole or in part. In the foregoing descriptions of the various embodiments, those embodiments which refer to another embodiment may be appropriately combined with one or more other embodiments, as will be appreciated by one of skill in the art. Furthermore, those of ordinary skill in the art will appreciate that the foregoing description is by way of example only, and is not intended to limit the invention.

Claims

1. A process for producing ethanol, comprising: hydrogenating acetic acid in a reactor in the presence of a catalyst to form a crude ethanol product comprising ethanol, ethyl acetate, water, and acetic acid;separating at least a portion of the crude ethanol product in a first distillation column to yield a first residue comprising acetic acid and a first distillate comprising ethanol, ethyl acetate, and water;splitting the first distillate into a main stream and a bypass stream;removing water from at least a portion of the main stream to yield an ethanol mixture stream comprising less than 10 wt. % water; andfeeding the ethanol mixture stream and at least a portion of the bypass stream to a second distillation column, wherein the water concentration of the feed to the second distillation column is from 0.5 wt. % to 15 wt. %.

2. The process of claim 1, wherein water is removed from the main stream using an adsorption unit, membrane, extractive column distillation, molecular sieve, or a combination thereof.

3. The process of claim 2, wherein the adsorption unit is a pressure swing adsorption unit or a thermal swing adsorption unit.

4. The process of claim 2, wherein the adsorption unit comprises molecular sieves.

5. The process of claim 1, wherein the ethanol mixture stream comprises less than 6 wt. % water.

6. The process of claim 1, wherein the ethanol mixture stream comprises less than 4 wt. % water.

7. The process of claim 1, wherein the first distillate is not condensed prior to removing the water.

8. The process of claim 1, further comprising condensing at least a portion of the first distillate and dehydrating the condensed portion of the first distillate.

9. The process of claim 1, wherein at least a portion of the water removed from the first distillate is returned to the first column.

10. The process of claim 1, wherein the ethanol mixture stream comprises ethanol in an amount from 20 to 95 wt. %, water in an amount from 0.01 to 7 wt. %, ethyl acetate in an amount from 0.01 to 60 wt. %, and acetaldehyde in an amount from 0.01 to 10 wt. %.

11. The process of claim 1, wherein the first residue further comprises water.

12. The process of claim 11, further comprising recovering acetic acid from the first residue and returning at least a portion of the recovered acetic acid to the reactor.

13. The process of claim 1, further comprising separating a portion of the feed to the second distillation column to yield a second residue comprising ethanol and a second distillate comprising ethyl acetate.

14. The process of claim 13, wherein the second residue comprises ethanol in an amount from 80 to 99.5 wt. %.

15. The process of claim 13, further comprising reducing the water content of the second residue to yield an ethanol product stream with reduced water content.

16. The process of claim 13, wherein at least a portion of the second distillate is recycled to the reactor.

17. The process of claim 1, wherein the acetic acid is formed from methanol and carbon monoxide, wherein each of the methanol, the carbon monoxide, and hydrogen for the hydrogenating step is derived from syngas, and wherein the syngas is derived from a carbon source selected from the group consisting of natural gas, oil, petroleum, coal, biomass, and combinations thereof.