RU2352551C2 - Obtaining branched aliphatic alcohols using joined technological flows of hydration installation and installation for dehydration-isomerisation - Google Patents

Abstract

FIELD: chemistry.

SUBSTANCE: method includes introduction of hydrocarbon Fisher-Tropsch flow, containing olefins and paraffins, said hydrocarbon Fisher-Tropsch flow containing from 5 to 80% of olefins, which have average number of carbon atoms from 10 to 17, and paraffins, into installation of hydration, where hydration installation is made in such way as to hydrate at least part of olefins in hydrocarbon Fisher-Tropsch flow to paraffins, and where at least part of components of hydrocarbon Fisher-Tropsch flow, which did not react, and at least part of hydrated olefins form second hydrocarbon flow; introduction of second hydrocarbon flow into installation of dehydration-isomerisation, where installation of dehydration-isomerisation is made in such way as to hydrate at least part of paraffins in second hydrocarbon flow to olefins, and where installation of dehydration-isomerisation is also made in such way as to isomerize at least part of linear olefins to branched olefins in presence of dehydration-isomerization catalyst, which contains hydrogen form of zeolite, having isotopic lattice structure of ferrierite, and where period duration is such that transformation of paraffins into olefins is lower than 40% mol, and where at least part of second hydrocarbon flow components that did not react, and at least part of products of dehydration and isomerisation reactions form third hydrocarbon flow, and third hydrocarbon flow contains olefins and paraffins, and where at least part of olefins in third hydrocarbon flow represent branched olefins; and introduction of at least part of third hydrocarbon flow into installation of hydroformilation, where installation of hydroformilation is made in such way as to hydroformilate at least part of olefins in third hydrocarbon flow obtaining aliphatic alcohols with average number of carbon atoms from 11 to 18, and where at least part of obtained aliphatic alcohols contain branched alkyl group.

EFFECT: reduction of expenditure.

18 cl, 7 tbl, 5 dwg, 6 ex

Description

BACKGROUND OF THE INVENTION

Cross reference to related application

This application claims the benefit of provisional application for US patent registration No. 60/511562, pending review October 15, 2003.

Technical field

This application generally relates to systems and methods for producing aliphatic alcohols. In particular, the embodiments described herein relate to systems and methods for producing branched aliphatic alcohols using hydrogenation and dehydrogenation-isomerization units.

Description of the prior art

Aliphatic alcohols are important compounds that can be used for various purposes or can be converted into other chemical compounds (for example, surfactants, sulfates). Surfactants can be used for various purposes (for example, as detergents, soaps, in the regeneration of oils).

The structure of an aliphatic alcohol may affect the properties of the surfactant and / or detergent produced from the aliphatic alcohol (for example, solubility in water, biodegradability and washing ability in cold water). For example, solubility in water can be affected by the linearity of the aliphatic portion of the aliphatic alcohol. With an increase in the linearity of the aliphatic portion, the hydrophilicity (i.e., affinity for water) of a surfactant based on aliphatic alcohol may decrease. Thus, the solubility in water and / or the washing properties of a surfactant based on aliphatic alcohol can be reduced. The incorporation of branches into the aliphatic portion of the alcohol-based surfactant may increase the solubility in cold water and / or the washing ability of the alcohol-based surfactant. The biodegradability of aliphatic alcohol surfactants, however, may decrease if the branches in the aliphatic portion of the alcohol surfactant include a large number of quaternary carbon atoms. The introduction of branches with a minimum number of quaternary carbon atoms into the aliphatic part of an aliphatic alcohol surfactant can increase the solubility in cold water and / or the washing ability of alcohol surfactants while maintaining the biodegradability of detergents.

The aliphatic portion of the aliphatic alcohol used to produce the surfactant may include one or more aliphatic alkyl groups as branches. Aliphatic alkyl groups which may form branches in the aliphatic part may be methyl, ethyl, propyl or higher alkyl groups. Quaternary or tertiary carbon atoms may be present if the aliphatic part is branched. The number of quaternary and tertiary carbon atoms may be the result of a branching scheme in the aliphatic part. In the present invention, the expression “aliphatic quaternary carbon atom” refers to a carbon atom that is not bound to any hydrogen atom.

Methods for preparing branched primary alcohol compositions are described in US Pat. No. 5,849,960 (Singleton et al.), Entitled “Highly Branched Primary Alcohol Compositions and Biodegradable Detergents Made Therefrom,” as well as US Pat. No. 6,150,322 (Singleton et al.), Under the name “Highly Branched Primary Alcohol Compositions and Biodegradable Detergents Made Therefrom.”

The essence of the invention

In one embodiment, aliphatic alcohols can be prepared by methods that include hydrogenation of olefins to paraffins. The feed stream to the hydrogenation unit may include linear olefins and paraffins with an average carbon number of 7 to 18. In one embodiment, the feed stream to the hydrogenation unit includes linear olefins and paraffins with an average carbon number 10 to 17. In the present invention, the expression “carbon number or number of carbon atoms” refers to the total number of carbon atoms in a molecule. The process stream of feed to the hydrogenation unit, in some embodiments, is obtained by the Fischer-Tropsch process.

At the hydrogenation unit, at least a portion of the olefins in the feed stream can be hydrogenated to produce paraffins. The resulting paraffin feed stream may be fed to a dehydrogenation-isomerization unit. At least a portion of the paraffins in the feed stream may be dehydrogenated to form olefins. The dehydrogenation-isomerization unit can also isomerize at least a portion of the obtained olefins and at least a portion of the olefins that are already present in the feed stream. The olefins obtained from the dehydrogenation-isomerization unit can be hydroformylated to produce aliphatic alcohols. At least a portion of aliphatic alcohols may have a branched aliphatic structure.

At least a portion of the paraffins in the feed stream may be dehydrogenated to produce olefins in a dehydrogenation-isomerization unit. At least a portion of the obtained olefins and at least a portion of the olefins that are already present in the feed stream can also be isomerized in a dehydrogenation-isomerization unit. The isomerization process converts linear olefins (e.g., unbranched olefins) to branched. Isomerized olefins can be hydroformylated to give the corresponding aliphatic alcohols. After hydroformylation of olefins, unreacted components can be separated from the obtained aliphatic alcohols. The paraffins and unreacted olefins in the separated stream can be recycled back to the dehydrogenation-isomerization unit.

The process conditions in the dehydrogenation-isomerization unit may be such that the obtained branched olefins have an average number of branches per olefin molecule from 0.7 to 2.5. Branched olefins may be, but are not limited to, methyl and / or ethyl branched olefins. The isomerization process can produce branched olefins that contain less than 0.5% quaternary aliphatic carbon atoms. The dehydrogenation-isomerization unit may contain a catalyst that has two functions: dehydrogenation of paraffins to olefins and isomerization of olefins into branched olefins.

In one embodiment, the dehydrogenation-isomerization unit may include a large number of zones, namely the first reaction zone and the second reaction zone. The first reaction zone may be a dehydrogenation zone. The second reaction zone may be an isomerization zone. A hydrocarbon stream consisting of olefins and paraffins may enter the dehydrogenation zone. At least a portion of the paraffins in the hydrocarbon stream may be dehydrogenated to olefins to produce an enriched stream. The olefin-rich stream may enter the isomerization zone. In the isomerization zone, at least a portion of the olefins from the olefin-rich stream may be isomerized to branched olefins. Branched olefins can be converted to aliphatic alcohols by hydroformylation. After hydroformylation of olefins, the stream of paraffins and unreacted olefins can be separated from the resulting aliphatic alcohol products. The paraffin and unreacted olefin stream may be recycled by directing at least a portion of the paraffin and unreacted olefin stream to the dehydrogenation-isomerization unit and / or to the stream entering the dehydrogenation-isomerization unit.

In one embodiment of the invention, the dehydrogenation-isomerization unit may have a composite (folded) layer configuration, which may include a dehydrogenation catalyst and a hydrocarbon stream isomerization catalyst. At least a portion of the paraffins in the hydrocarbon stream may be dehydrogenated to olefins to produce an enriched stream. The olefin-rich stream may enter the isomerization zone. In the isomerization zone, at least a portion of the olefins in the olefin-rich stream can be isomerized to branched olefins, which can be converted to aliphatic alcohols by hydroformylation. After hydroformylation of olefins, the stream of paraffins and unreacted olefins can be separated from the resulting aliphatic alcohol products. The stream of paraffins and unreacted olefins can be recycled by directing at least a portion of the stream of paraffins and unreacted olefins back to the dehydrogenation-isomerization unit and / or to the stream entering the dehydrogenation-isomerization unit.

In some embodiments, at least a portion of the aliphatic alcohols may be sulfated to form aliphatic sulfates. In some embodiments, aliphatic sulfates may contain branched alkyl groups.

In some embodiments of the invention, at least a portion of the resulting aliphatic alcohols can be oxyalkylated to produce oxyalkyl alcohols. In some embodiments, at least a portion of the resulting aliphatic alcohols may contain branched alkyl groups. In some embodiments of the invention, at least a portion of the branched aliphatic alcohols obtained can be ethoxylated to form branched ethoxyalkyl alcohols. At least a portion of the hydroxyalkyl alcohols may be sulfated to form hydroxyalkyl sulfates. In some embodiments, the hydroxyalkyl sulfates may contain branched alkyl groups.

Brief Description of the Drawings

The advantages of the present invention will be apparent to a person skilled in the art when considering the following detailed description of embodiments of the invention and when considering the accompanying drawings, where:

FIG. 1 shows a schematic diagram of an embodiment of a system for the production of aliphatic alcohols using a hydrogenation unit and a dehydrogenation-isomerization unit;

FIG. 2 shows a schematic diagram of an embodiment of a system for the production of branched aliphatic alcohols using a hydrogenation unit, a dehydrogenation-isomerization unit, and a separation unit for separating branched olefins from linear olefins and paraffins;

FIGA-B show the schematic diagrams of embodiments of a system for the production of branched aliphatic alcohols using a hydrogenation unit and a dehydrogenation-isomerization unit with two zones;

FIG. 3 shows a schematic diagram of an embodiment of a system for the production of branched aliphatic alcohols using a hydrogenation unit and a dehydrogenation-isomerization unit with a catalyst configuration in the form of a composite layer.

Although the present invention is subject to various modifications and alternative forms, specific embodiments of the invention are presented by way of example in the drawings and will be discussed in detail. It should be understood that the drawings and detailed description are not intended to limit the invention to the particular form described, but rather, the invention should cover all modifications, equivalents, and alternatives that fall within the spirit and scope of the present invention as defined in the appended claims.

Detailed Description of Embodiments

Hydrocarbon products can be synthesized from synthesis gas (i.e. a mixture of hydrogen and carbon monoxide) using the Fischer-Tropsch process. Synthesis gas can be obtained by partial combustion of oil (e.g., coal, hydrocarbons), conversion of natural gas, or due to its partial oxidation. The Fischer-Tropsch process catalytically converts synthesis gas into a mixture of products that include saturated hydrocarbons, unsaturated hydrocarbons and a small amount of oxygen-containing products. The products of the Fischer-Tropsch process can be used to produce fuels (e.g. gasoline, diesel), lubricants and waxes.

Process streams of Fischer-Tropsch synthesis can also be used to obtain consumer products that are of economic importance. For example, linear olefins are general-purpose products that can be used to produce surfactants. The use of part of the process stream to obtain linear olefins can increase its efficiency.

Surfactants derived from branched olefins may have other properties than surfactants derived from linear olefins. For example, surfactants derived from branched olefins have increased solubility in water and / or improved detergent properties compared to surfactants derived from linear olefins. The biodegradability of a surfactant, however, can be affected by the presence of quaternary carbon atoms in its branched part. Surfactants derived from branched olefins with a minimum number of Quaternary carbon atoms may have similar biodegradability as surfactants derived from linear olefins. The production of branched olefins from the Fischer-Tropsch process stream can increase its economic value. In some embodiments, linear olefins can be converted to branched olefins with a minimum number of quaternary carbon atoms using an isomerization catalyst. An increase in the number of branched olefins obtained from the Fischer-Tropsch process stream can increase its economic value.

Methods are described for increasing the number of branched olefins obtained from a process stream that includes a certain amount of olefins, which as a result increases its economic value. Such methods can be used for both Fischer-Tropsch process streams and product streams from other sources that contain hydrocarbons.

The composition of the hydrocarbon feed stream may include paraffins and olefins. At least a portion of the hydrocarbon stream may consist of linear paraffins and olefins containing at least 4 to 10 carbon atoms. The hydrocarbon feed stream can be obtained by Fischer-Tropsch processes or ethylene oligomerization. Fischer-Tropsch catalysts and reaction conditions can be selected so as to obtain a specific mixture of products in the stream. For example, the Fischer-Tropsch catalyst and reaction conditions can be selected so as to increase the number of olefins and reduce the amount of paraffins and oxidized compounds in the stream. On the other hand, the catalyst and reaction conditions can be selected so as to increase the amount of paraffins and reduce the amount of olefins and oxidized compounds in the stream.

The catalyst used in the Fischer-Tropsch process may be compounds of Mo, W, Group VIII elements, or a combination thereof. Group VIII compounds include, but are not limited to, iron, cobalt, ruthenium, rhodium, platinum, palladium, iridium, and osmium compounds. Combinations of Mo, W compounds and Group VIII elements can be obtained in the form of free metals. In an embodiment of the invention, combinations of Mo, W compounds and Group VIII elements can be obtained as alloys. Combinations of Mo, W compounds and Group VIII elements may be obtained in certain embodiments as oxides, carbides or other compounds. In other embodiments, combinations of Mo, W compounds, and Group VIII elements may be prepared as salts. Iron and cobalt based catalysts are used in industry as Fischer-Tropsch catalysts. Catalysts of ruthenium under high pressure conditions favor the preferred formation of high molecular weight paraffins with a high melting point. Fischer-Tropsch synthetic catalysts may include fused iron. In some embodiments of the invention, the Fischer-Tropsch fused iron catalyst may include a promoter (e.g., potassium or oxides on a silica support, an alumina support or an aluminosilicate support). In the Fischer-Tropsch catalyst, cobalt metal may be used. With the appropriate choice of the substrate, promoters and combinations of other metals, the cobalt-based catalytic system can be oriented to the production of a composition enriched with the required hydrocarbons. It is known that catalysts based on an iron-cobalt alloy are selective for the production of olefins. Catalysts and combinations for the production of various hydrocarbons using the Fischer-Tropsch process are generally known.

Although a stream of the Fischer-Tropsch process is mentioned, any stream of olefins and saturated hydrocarbons may be used. A large number of Fischer-Tropsch streams may contain from 5 to 80% olefins, with the remaining amount being saturated hydrocarbons, including paraffins and other compounds.

In some embodiments, feed streams containing olefins and paraffins can be obtained through cracking of high molecular weight paraffins or oligomerization of olefins. Industrial-derived olefins produced by ethylene oligomerization are sold in the United States by the Chevron Phillips Chemical Company, Shell Chemical Company (as NEODENE ^® ) and British Petroleum. The cracking of high molecular weight paraffins to produce α-olefins and paraffin feeds is described in US Pat. No. 4,579,986 (Sie) under the name “Process for The Preparation of Hydrocarbons” and in US Patent Application No. 10/153955 (Ansorge et al.) Under the name “Process for the Preparation of Linear Olefins and Use Thereof to Prepare Linear Alcohols ”. Specific processes for preparing linear olefins from ethylene are described in US Pat. No. 3,676,523 (Mason) under the name “Alpha-Olefin Production”, US Patent No. 3,686,351 (Mason) under the name “Alpha-Olefin Production”; in US patent No. 3737475 (Mason) under the name “Alpha-Olefin Production” and in US patent No. 4020121 (Kister et al.) under the name “Oligomerization Reaction System”. Most of the processes mentioned above produce α-olefins. Higher linear internal olefins can be obtained industrially (for example, by chlorination-dehydrochlorination of paraffins, dehydrogenation of paraffins, isomerization of α-olefins).

In one embodiment, a feed stream is prepared to produce a hydrocarbon stream that contains branched olefins. Such branched olefins can be converted to branched aliphatic alcohols using various methods. The feed stream may contain paraffins from 50 to 90 wt.% Based on the feed stream. In some embodiments, the feed stream may contain more than 90 wt.% Paraffins. The feed stream may also include olefins. The olefin content in the feed stream may be from 10 to 50 wt.%. In some embodiments, the feed stream may contain more than 90 wt.% Olefins.

The composition of the feed stream may include hydrocarbons with an average number of carbon atoms from 4 to 30. In some embodiments, the average number of carbon atoms in a feed stream can be from 4 to 24. In other embodiments, the average number of carbon atoms in a feed stream can be from 4 to 18. The average average number of carbon atoms in a feed stream can be from 7 to 18. In some embodiments, the average number of carbon atoms in a feed stream can be from 10 to 17. In some embodiments of the invention, the average number of carbon atoms in the feed stream can be from 10 to 13. In other embodiments, the average number of carbon atoms in the feed stream can be from 14 to 17.

Branched chain olefins can be converted to branched aliphatic alcohols (e.g., branched primary alcohols) by hydroformylation. The term “hydroformylation”, as used herein, refers to the production of alcohols from olefins through a carbonylation and hydrogenation process. Other processes may be used to produce aliphatic alcohols from olefins. Examples of other processes for producing aliphatic alcohols from olefins include, but are not limited to, hydration, oxidation and hydrolysis, sulfation and hydration, and epoxidation and hydration. The composition of the alcohol product stream may include aliphatic alcohols having an average number of carbon atoms from 5 to 31. In an embodiment of the invention, the average number of carbon atoms of aliphatic alcohols in the alcohol product stream can be from 7 to 20. In some embodiments, the average number of carbon atoms is aliphatic alcohols in the alcohol product stream may range from 11 to 18. In some embodiments, the average number of carbon atoms of aliphatic alcohols in the alcohol product stream m may be from 11 to 14. In other embodiments, the average number of carbon atoms of aliphatic alcohols in the alcohol product stream may be from 15 to 18.

In some embodiments of the invention, the first hydrocarbon stream may contain undesired compounds (e.g., oxygenates and dienes) that can reduce the selectivity of the catalyst in the process used to produce aliphatic alcohols. Removal of undesired compounds can be carried out by hydrogenation of the first hydrocarbon stream. Hydrogenation of the first hydrocarbon stream in some embodiments of the invention may produce a hydrocarbon stream containing more than 90 wt.% Paraffins. The hydrogenated hydrocarbon stream may be dehydrogenated to produce an olefin stream. The catalyst used in the dehydrogenation process can determine the position of the double bond. In some embodiments of the invention, the olefin hydrocarbon stream may comprise olefins, of which α-olefins with a linear carbon skeleton structure comprise more than 70 wt.%. In other embodiments, the olefin hydrocarbon stream may include olefins, of which more than 50% are internal olefins.

The first hydrocarbon stream may be introduced into the hydrogenation unit 110 through the first pipe 112, as shown for the system 100 in FIG. 1. The first hydrocarbon stream includes olefins and paraffins. In the hydrogenation unit 110, at least a portion of the olefins in the first hydrocarbon stream may be hydrogenated to paraffins to form a second hydrocarbon stream.

The reaction conditions in the hydrogenation unit 110 can be controlled to hydrogenate olefins and dienes and remove oxygenates. The operating temperature of the hydrogenation unit 110 may be from 100 to 300 ° C. In some embodiments, the operating temperature may be from 150 to 275 ° C. In other embodiments, the implementation of the operating temperature may be from 175 to 250 ° C. Working pressure can be from 5 to 150 atmospheres (from 506 to 1520 kPa). In some embodiments, the operating pressure may be from 10 to 500 atmospheres (from 1013 to 5065 kPa).

Hydrogenation processes can be carried out using any type of catalyst bed arrangement (e.g., fluidized bed, moving bed, suspended phase bed or fixed bed). In some embodiments, a fixed bed may be used. In a fixed-bed system, hydrogen can be fed to the hydrogenation stage with a gas hourly space velocity of 100 normal liters of gas / liter of catalyst / hour (nL / L / hour) to 1000 nL / L / hour. In some embodiments, hydrogen may be supplied at a gas hourly space velocity of 250 to 5000 nL / L / h. The definition of "hourly space velocity of gas, expressed in units of normal liter / liter of catalyst / hour" used in the description, means the volume of gas in liters under standard conditions of 0 ° C and a pressure of 760 mm Hg.

A wide range of hydrogenation catalyst compositions are generally known and commercially available. In some embodiments, the hydrogenation catalyst may contain one or more metals from groups VIB and VII of the Periodic Table of the Elements. In some embodiments, metals include, but are not limited to, molybdenum, tungsten, cobalt, nickel, ruthenium, iridium, osmium, platinum and palladium. The hydrogenation catalyst may contain a refractory oxide or silicate as a binder.

Hydrogenation reaction conditions and catalysts are described in European patent No. 0583836 (Eilers et al.) Under the name “Process for Preparation of Hydrocarbon Fuels”, European patent No. 0668342 (Eilers et al.) Under the name “Lubricating Base Oil Preparation Process”. Hydrogenation reaction conditions and catalysts are also described in US Pat. No. 5,371,308 (Gosselink et al.), Entitled “Process for Preparation of Lower Olefins”.

In some embodiments, a second hydrocarbon stream that contains paraffins and olefins may be introduced to the dehydrogenation-isomerization unit. A dehydrogenation-isomerization unit can replace two independent units (for example, an isomerization unit and a dehydrogenation unit). The dehydrogenation-isomerization unit can dehydrogenate paraffins to olefins and isomerize the resulting olefins and / or starting olefins present in the hydrocarbon stream to branched olefins. In an embodiment of the invention, the catalyst can carry out dehydrogenation-isomerization of hydrocarbons in a second hydrocarbon stream. In some embodiments, the catalyst may be one. In some embodiments, the implementation of the catalyst may be a mixture of two catalysts (for example, a dehydrogenation catalyst and an isomerization catalyst). In other embodiments, two separate catalysts located in different zones or in a composite layer configuration in a single dehydrogenation-isomerization unit can carry out a dehydrogenation-isomerization process. In this case, the definition of “dehydrogenation-isomerization catalyst” may be one or more catalysts.

In some embodiments of the invention, the dehydrogenation-isomerization unit may have several entry points to allow for the supply of various process streams. Process streams can be obtained from other process plants and / or from storage facilities. Examples of process streams include, but are not limited to, a dilute hydrocarbon stream and / or other hydrocarbon streams that include olefins and paraffins derived from other processes. Used in this case, the definition of "input to the installation of dehydrogenation-isomerization" refers to the input of process flows to the installation of dehydrogenation-isomerization through one or more entry points.

A second hydrocarbon stream including a mixture of olefins and paraffins can be introduced to the dehydrogenation-isomerization unit 114 through a second line 116. In the dehydrogenation-isomerization unit 114, at least a portion of the paraffins in the second hydrocarbon stream can be dehydrogenated to olefins. At least a portion of the obtained olefins and at least a portion of the olefins that were already present in the feed stream can be isomerized to produce a third hydrocarbon stream. The isomerization process converts linear olefins (i.e., unbranched olefins) to branched olefins.

The catalyst used for the dehydrogenation-isomerization of the second hydrocarbon stream can be obtained on the basis of a zeolite catalyst modified with one or more metals or their compounds. The catalyst used in the dehydrogenation-isomerization unit 114 for processing olefins in a second hydrocarbon stream can be effective for skeletal isomerization of linear olefins in a process stream into olefins having an average number of branches per chain of olefin molecule of more than 0.7. In some embodiments, the average number of branches per chain of an olefin molecule is from 0.7 to 2.5. In some embodiments, the average number of branches per chain of an olefin molecule is from 1.0 to 2.2.

The dehydrogenation-isomerization catalyst may comprise a zeolite having at least one channel with a crystallographically free diameter greater than 4.2 Å, or less than 7 Å, measured at room temperature. In this case, the definition of “channel diameter or size” refers to the channel diameter or size effective for diffusion. Zeolite may not have channels with a free diameter greater than 7 Å. The catalyst may contain at least one channel having a crystallographically free inlet diameter greater than 4.2 Å, or less than 7 Å. The catalyst may not have a channel with an inlet diameter that exceeds 7 Å of the upper boundary of the interval. Zeolites with channel diameters greater than 7 Å may be sensitive to the formation of undesired olefin by-products (e.g., aromatization, oligomerization, alkylation, coking). In some embodiments, the zeolite may not contain a channel having a free diameter along either the x or y planes greater than 4.2 Å. The small size of the channels can impede diffusion of branched olefins and / or from pore channels. A zeolite may have at least one channel with a free diameter greater than 4.2 E and less than 7 Å.

In an embodiment of the invention, the olefin molecule due to the high length of the carbon chain may not enter the zeolite channel, diffuse through it and not exit through the other end of the channel. The degree of branching detected when the olefin passes through the zeolite may not correspond to the theoretical one, provided that each olefin molecule has passed through the channels. Most olefins can partially penetrate the channel to a distance effective for branching part of the chain within the channel, and then exit the channel after isomerization. In one embodiment of the method, in order to obtain aliphatic alcohols, the olefin molecules in the hydrocarbon stream may advantageously have a structure branched along the edges of the main carbon chain and is essentially linear with respect to the center of the molecule (for example, at least 25% of the carbon atoms in the central part are unbranched).

In some embodiments, the zeolite catalyst structure may comprise channels having free diameters greater than 4.2 Å and less than 7 Å along both x and y planes in the form [001]. Zeolites with specific channel sizes can be called zeolites with middle or intermediate channels and usually have a 10-T element (or composite 12-T element) of the annular channel structure in one species and 9-T element or less (small pore) in another form , if he is. In zeolite there are no restrictions on the number of channels or on their orientation (for example, parallel, mutually unconnected intersections or interconnected at any angle).

Examples of zeolites with channel sizes from 4.2 to 7 Å are molecular sieves, ferrierite, A1RO-31, SAPO-11, SAPO-31, SAPO-41, FU-9, NU-10, NU-23, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-50, ZSM-57, SUZ-4A, MeAPO-11, MeAPO-31, MeAPO-41, MeAPSO-31 and MeAPSO-41, MeAPSO- 46, ELAPO-11, ELAPO-31, ELAPO-41, ELAPO-11, ELAPO-31 and ELAPO-41, laumontite, cancrinite, offretite, hydrogen form of stilbite, magnesium or calcium form of mordenite and partheitis. Isotypic structures of zeolite lattices, known under other names, can also be considered equivalent. The zeolite lattice described by Flanigen et al. in the publication “Aluminophospate Molecular Sieves and the Periodic Table”, New Development in Zeolite Science Technology, 1986, Kodansha Ltd., Tokyo, Japan.

A large number of natural zeolites, such as ferrierite, heilandite, and stilbit, can be similar to one-dimensional porous structures with pore sizes of about or slightly less than 4.2 Å in diameter. US Pat. No. 4,795,623 (Evans) entitled “Time Effective Method for Preparing Ferrierite” and US Pat. No. 4,942,027 (Evans) called “Method for Preparing Ferrierite” describe the conversion of channels in natural zeolites to larger channels. Channels in natural zeolites can be converted into zeolites with the desired larger channels by removal of the bound alkali metal or alkaline earth metal by known methods (e.g., by exchanging with ammonium ion, optionally followed by calcination, to obtain the zeolite in substantially hydrogen form). Replacing the bonded alkaline or alkaline earth metal with a hydrogen form, you can increase the diameter of the channel. In some embodiments, natural zeolites (e.g., some forms of mordenite) may have a channel size greater than 7 Å. The channel size can be reduced by replacing the alkali metal with larger ions (for example, a larger bulk alkaline earth metal).

In some embodiments, zeolites may have an isotypic (or homeotypic) ferrierite structure. Parallel channels in the aluminosilicate lattice can be significant structural features of ferrierite detected by X-ray crystallography. Parallel channels may have an elliptical cross section. Zeolites having an isotypic lattice structure of ferrierite are described in European patent No. 55529 (Seddon et al.) Under the name “Zeolites” and in European patent No. 103981 (Whittam) under the name “Zeolites”. Zeolites having an isotypic ferrierite lattice are also described in US Pat. No. 4,016,245 (Plank et al.), US Pat. No. 4,578,259 (Morimoto et al.) Under the names “Process for Preparing a Crystalline Aluminosilicate”, “Crystalline Zeolite and Method of Preparing Same "And in US patent No. 4375573 (Young et al.) Entitled" Selective Production and Reaction of P-Disubstituted Aromatics Over Zeolite ZSM-48 ".

In an embodiment of the invention, the hydrogen form of ferrierite (H-ferrierite) can be considered essentially as one-dimensional. N-Ferrierite may have parallel working channels. H-Ferrierite can have elliptical channels, which have free diameters from 4.2 to 5.4 Å along the x and y planes in the form [001]. The channels may be large enough to allow linear olefins to enter and diffuse methyl branched isoolefins through the channel. The channels can be small enough to slow down the formation of coke deposits. Methods for producing various H-ferrierites are described in US Pat. No. 5,985,238 (Pasquale et al.) Under the name “Process for Preparing Ferrierite”, in US Pat. No. 4,251,499 (Nanne et al.) Under the name “Process for Preparation of Ferrierite” in the patent. US No. 4795623 (Evans) under the name “Time Effective Methods for Preparing Ferrierite” and in US patent No. 4942027 (Evans) under the name “Methods for Preparing Ferrierite”.

In some embodiments, the dehydrogenation-isomerization catalyst may be mixed with a refractory oxide that serves as a binder. Suitable refractory oxides include, but are not limited to, natural clays (e.g., bentonite, montmorillonite, attapulgite and kaolin), alumina, silica, aluminosilicate, hydrated alumina, titanium oxide, zirconia, or mixtures thereof.

Examples of alumina-based binders include, but are not limited to, pseudoboehmite, gamma alumina, and bayerite. Binders based on aluminum oxide may be commercially available (e.g., LaPoche Chemicals manufactures VERSAL ^® aluminas and Sasol manufactures CATAPAL ^® alumina). In an embodiment of the invention, finely divided alumina powders can be used as alumina binders when extrusion is used to produce the catalyst. Fine alumina powders can have a degree of dispersion of more than 50% in an aqueous acid dispersion having an acid content of 0.4 milligram equivalent acid (acetic) per gram of powder. Such high-dispersity aluminas may be represented example CATAPAL ^® alumina manufactured by Sasol.

The mass ratio of zeolite to a binder material can be from 10:90 to 99.5: 0.5. In some embodiments of the invention, the mass ratio may be from 75:25 to 99: 1. In other embodiments of the invention, the mass ratio of zeolite to binder material may be from 80:20 to 98: 2. In some embodiments, the weight ratio of zeolite to binder may be from 85:15 to 95: 5 based on the anhydrous base.

In some embodiments of the invention, a dehydrogenation-isomerization catalyst may be prepared with one or more monocarboxylic acids and / or inorganic acids. In addition to monocarboxylic and / or inorganic acids, at least one organic acid with at least two carboxyl groups (polycarboxylic acid) can be used. Monocarboxylic acids may have substituted or unsubstituted hydrocarbon groups containing from 1 to 20 carbon atoms. Hydrocarbon groups may be aliphatic, cyclic or aromatic. Examples of monocarboxylic acids containing from 1 to 20 carbon atoms include, but are not limited to, acetic acid, formic acid, propionic acid, butyric acid, caproic acid, glycolic acid, lactic acid, hydroxybutyric acid, hydroxycyclopentanoic acid, salicylic acid, mandelic acid, benzoic acid and fatty acids. Examples of inorganic acids include, but are not limited to, nitric acid, phosphoric acid, sulfuric acid, and hydrochloric acid.

The polycarboxylic acid in some embodiments of the invention may be an organic acid with two or more carboxyl groups attached via a carbon-carbon bond to a hydrocarbon segment. The compound may be in any position of the hydrocarbon segment. A polycarboxylic acid may have a hydrocarbon segment with less than 10 carbon atoms. The hydrocarbon segment may be aliphatic, cyclic or aromatic. The hydrocarbon segment may not have carbon atoms in the case of oxalic acid with two carboxyl groups attached via a carbon-carbon bond. Examples of polycarboxylic acids include, but are not limited to, tartaric acid, citric acid, maleic acid, oxalic acid, adipic acid, malonic acid, galactaric acid, 1,2-cyclopentane-dicarboxylic acid, maleic acid, fumaric acid, itaconic acid, phthalic acid, terephthalic acid, phenylmalonic acid, hydroxyphthalic acid, dihydroxyfumaric acid, tricarballylic acid, benzene-1,3,5-tricarboxylic acid, isolimonic acid, mucus acid and glucaric acid ota. Polycarboxylic acids can be any of the isomers of these acids. In some embodiments, the polycarboxylic acids may be any stereoisomers of the above acids. In one embodiment of the invention, polycarboxylic acids with at least two carboxyl groups and at least one hydroxyl group are used. In one embodiment of the invention, citric acid, tartaric acid and malic acid can be used as polycarboxylic acids.

The metals introduced into the dehydrogenation-isomerization catalyst may be metals that promote the oxidation of coke in the presence of oxygen at temperatures above 250 ° C and the dehydrogenation of paraffins. The term “metal (s)”, as used herein, refers to metals with a zero oxidation state and / or higher oxidation states (for example, metal oxides). As used herein, the term “coke” refers to the thermal decomposition product of larger molecules to smaller molecules.

The metals used in the dehydrogenation-isomerization catalyst can be transition and rare earth metals. Metals that accelerate the oxidation of coke are, but are not limited to, metals of groups IB, VB, VIB, VIIB, VIII of transition metals of the Periodic Table and / or combinations thereof. In some embodiments of the invention, Pd, Pt, Ni, Co, Mn, Ag, Cr and / or combinations thereof may be used in a dehydrogenation-isomerization catalyst. In other embodiments of the invention, metal oxides, such as, but not limited to, chromium oxide, iron oxide, noble metals and mixtures thereof, may be present in the catalyst as coke oxidizing compounds.

The amount of metal introduced can range from 5 parts per million (ppm) and up to 15 wt.%. In some embodiments, the amount of metal may be from 5 ppm to 10 wt.%. In some embodiments, the amount of metal may be from 5 ppm to 5 wt.%.

Noble metals (e.g., platinum and / or palladium) can be used in lower amounts than other metals incorporated into the zeolite and / or binder. The definition of "noble metals" used in this case refers to the metals of the group, which includes platinum, palladium, iridium, ruthenium, osmium and rhodium. In some embodiments, the amount of noble metal (s) may be from 5 ppm to 2 wt.% Metal, based on the final catalyst. In some embodiments, the amount of noble metal (s) may be from 5 to 1000 ppm of metal per final catalyst. In other embodiments, the amount of noble metal (s) may be from 5 to 3,000 ppm of metal based on the final catalyst. The amount of noble metal (s) used in the dehydrogenation-isomerization catalyst, in some embodiments, may be from 5 to 2000 wt. ppm metal based on the final catalyst. The amount of precious metal (s) sufficient to accelerate regeneration without compromising the properties of the catalyst may be from 30 to 100 ppm. Higher amounts of platinum and / or palladium (for example, greater than 2 wt.%) Can adversely affect the catalyst (for example, duration, olefin isomerization activity, selectivity).

In an embodiment of the invention, the zeolite powder and the alumina powder may be mixed (for example, formed into a paste) with water and one or more catalyst metal compounds. The resulting mixture can be molded into granules. The catalysts obtained by molding the paste can have excellent olefin isomerization properties compared to impregnated catalysts. The definition of “pasta molding” refers to the mixing of powders to which enough water has been added to produce a thick paste and where the mixing is carried out by cutting the paste. Commercially available paste machines such as the Lancaster Mix Muller and Simpson Mix Muller can be used commercially.

In some embodiments, granules may be extruded. One or more peptizing acids (e.g., nitric acid, acetic acid, citric acid, or mixtures thereof) can be added to the mixture, and optional extrusion additives such as cellulose derivatives (e.g. METHOCEL ^® F4M, hydroxypropyl methyl cellulose, manufactured by The Dow Chemical, can be used). Company). The amount of peptizing acid used to produce a plastic extrudable material can be determined using routine tests. The definition of "granules" used in the description refers to any configuration or form of the combined material.

In some embodiments, a noble metal, such as platinum and / or palladium, can be added to the zeolite catalyst after granules are obtained. Conventional metal introduction methods known to those skilled in the art (for example, impregnation, ion exchange of a noble metal and co-molding into a paste) can be used to produce working catalysts that are useful in the dehydrogenation-isomerization of paraffins. Adding noble metals to the catalyst may contribute to the dehydrogenation reaction of paraffins. Pellets containing precious metals can be calcined at temperatures from 250 to 700 ° C. In some embodiments of the invention, the calcination temperature may be from 300 to 600 ° C. In some embodiments of the invention, the calcination temperature may be from 450 to 525 ° C.

The dehydrogenation-isomerization catalyst can be brought into contact with the second hydrocarbon stream in the dehydrogenation-isomerization unit 114 under various conditions to dehydrogenate at least a portion of the paraffins to olefins, to at least partly produce the olefins obtained. In the dehydrogenation-isomerization unit 114, reaction temperatures can range from 300 to 700 ° C. In some embodiments, reaction temperatures may range from 350 to 550 ° C. The total pressure in the dehydrogenation-isomerization unit 114 during the reaction may be from 0.010 to 25.0 atmospheres (from 1 to 2534 kPa). In one embodiment of the invention, the total pressure in the dehydrogenation-isomerization unit 114 during the reaction may be from 0.010 to 15.0 atmospheres (from 1 to 1520 kPa). In one embodiment of the invention, the total pressure in the dehydrogenation-isomerization unit 114 during the reaction may be from 1 to 5.0 atmospheres (from 101 to 507 kPa). To prevent coking, hydrogen may be supplied together with the second hydrocarbon stream. Hydrogen gas and paraffins present in the second hydrocarbon stream may be supplied at a molar ratio of hydrogen gas to paraffin from 0.1 to 20. In some embodiments, the molar ratio of hydrogen gas to paraffin may be from 1 to 10.

The residence time in the dehydrogenation-isomerization unit 114 can be chosen so that the degree of conversion of paraffins to olefins remains below 40 mol%. In one embodiment of the invention, the degree of conversion is from 5 to 30 mol.%. By maintaining a low degree of conversion, side reactions (for example, diene formation and cyclization reactions) can be minimized. The degree of conversion of olefins can be increased by changing the reaction conditions (for example, temperature, residence time) as long as adverse reactions remain below acceptable limits. The olefins obtained in the dehydrogenation-isomerization unit 114 may have higher degrees of branching than the paraffin fed to the dehydrogenation-isomerization unit. It should be understood that the concentration of olefins obtained through the dehydrogenation-isomerization unit 114 may be limited by the thermodynamic equilibrium of olefins and paraffins at a given reaction temperature. The conditions for isomerization of olefins in the dehydrogenation-isomerization unit 114 can be changed so that the number of carbon atoms in the olefins before and after isomerization is substantially the same.

The branched olefins obtained in the dehydrogenation-isomerization unit 114 may contain methyl, ethyl and / or longer carbon chains as branches. Nuclear magnetic resonance spectroscopy on protons ( ¹ H-NMR) of a composition of isomerized olefins can be performed. Branched olefins may contain quaternary and / or tertiary aliphatic carbon atoms. In some embodiments of the invention, the amount of aliphatic carbon atoms obtained in the dehydrogenation-isomerization unit can be minimized. ¹ H-NMR spectra of olefins can show the degree of olefin isomerization in the hydrocarbon stream. ¹ H-NMR spectroscopy is able to identify a wide range of olefin structures.

In one embodiment, the average number of branches per olefin molecule present in the resulting composition of branched olefins may be more than 0.7. In some embodiments, the average number of branches per olefin molecule present in the resulting composition of branched olefins is from 0.7 to 2.5. In some embodiments, the average number of branches per olefin molecule present in the resulting composition of branched olefins is from 0.7 to 2.2. In some embodiments, the average number of branches per olefin molecule present in the resulting composition of branched olefins is from 1.0 to 2.2. The degree of branching in the product can be controlled by changing the process conditions at the dehydrogenation-isomerization unit. For example, high temperatures and lower feed rates can result in a higher degree of branching. Methyl branches may comprise from 20 to 99% of their total number. In some embodiments, methyl branches may comprise more than 50% of the total number of branches present in olefin molecules. The number of ethyl branches in olefin molecules may in some embodiments be less than 30% of the total number of branches. In other embodiments of the invention, the number of ethyl branches, if present, may be from 0.1 to 2% of the total number of branches. Branches other than methyl and ethyl, if present, may comprise less than 10% of the total number of branches.

Aliphatic quaternary carbon atoms present in branched olefins may comprise less than 2% of the number of carbon atoms present. In one embodiment of the invention, the number of aliphatic quaternary carbon atoms present is less than 1% of the number of carbon atoms present. In an application in which biodegradability is important, the number of aliphatic quaternary carbon atoms may be less than 0.5% of the number of carbon atoms present. In one embodiment, the number of aliphatic quaternary carbon atoms is less than 0.3% of the number of carbon atoms present. In another embodiment, the number of aliphatic quaternary carbon atoms present in the branched olefin composition is from 0.01 to 0.3% of the number of carbon atoms present.

The third hydrocarbon stream may exit from the dehydrogenation-isomerization unit 114 and transferred to other process units (eg, hydroformylation unit, separation unit, alkylation unit) through the third pipeline 118. At least a portion of the third hydrocarbon stream may exit from the dehydrogenation-isomerization unit 114 and can be introduced to the hydroformylation unit 120 through a third conduit 118. In the hydroformylation unit 120, at least a portion of the olefins in the third carbohydrate native stream can be converted to olefins. At least a portion of the resulting alcohols and at least a portion of the unreacted components of the third hydrocarbon stream may form a hydroformylation reaction stream.

In one embodiment of the invention, olefins can be separated, if desired, from a hydrocarbon stream using known methods in the art (for example, by distillation, using molecular sieves, extraction, adsorption, absorption / desorption and / or using membranes). The separation of at least a portion of the branched olefins from linear olefins and paraffins can increase their concentration upon admission to the hydroformylation unit. In addition, the separation of at least a portion of the branched olefins from linear olefins and paraffins may affect the ratio of linear olefins to branched olefins obtained in a hydroformylation unit.

Referring to FIG. 2, a third hydrocarbon stream may exit the dehydrogenation-isomerization unit 114 and enter the separation unit 122 via a separation line 124. The separation unit 122 may produce at least two streams, a branched olefin stream and a linear olefin stream and paraffins. At separation unit 122, a third hydrocarbon stream may be contacted with organic and / or inorganic molecular sieves (e.g., zeolites or urea) with a selected pore size for branched olefins and / or linear olefins and paraffins. Subsequent desorption (e.g., solvent desorption) of at least a portion of the branched olefins and / or at least a portion of the linear olefins and paraffins from the molecular sieves may produce at least two streams (e.g., a branched olefin stream and a linear olefins and paraffins).

Separation unit 122 may include a fixed bed containing an adsorbent for separating a third hydrocarbon stream to produce a stream containing branched olefins and paraffins and a stream containing linear olefins and paraffins. The separation temperatures at separation unit 122 may range from 100 to 400 ° C. In some embodiments of the invention the separation temperature may be from 180 to 380 ° ^C. Separation of the separating device 122 can be conducted at a pressure of 2 to 7 atmospheres (from 202 to 710 kPa). In some embodiments, pretreatment of the third hydrocarbon stream may be performed to prevent adsorbent poisoning. An example of the adsorption / desorption process is a Molex process (Molex) using separation techniques Sorbex ^® (UOP process, UOP, Des Plaines, IL) . Adsorption / desorption processes are described in US Pat. No. 6,225,518 (Sohn et al.) Under the name “Olefinic Hydrocarbon Separation Process”, in US Pat. No. 5,292,990 (Kantner et al.) Under the name “Zeolite Compositions for use in olefinic separations” and in the patent US No. 5276246 (McCulloch et al.) Entitled “Process for Separating Normal Olefins from non-normal Olefins”.

At least a portion of the linear olefin and paraffin stream may be transferred to other process units or may be stored in place. In one embodiment of the invention, at least a portion of the linear olefin and paraffin stream may be mixed with the first hydrocarbon stream in conduit 112 through a linear olefin and paraffin recycle conduit 126. The combined stream may enter hydrogenation unit 110 through first conduit 112 to continue the process of obtaining aliphatic alcohols. In one embodiment, a linear olefin and paraffin stream may be introduced directly to hydrogenation unit 110.

At least a portion of the branched olefins stream may be transferred and used in other process streams and / or stored in place through the branched olefins pipeline 128. In some embodiments, at least a portion of the branched olefins stream may be sent to a separation unit 122 and may be introduced into the third conduit 118 through a branching olefin conduit 128. In another embodiment of the invention, at least a portion of the branched olefin stream may exit new separation 122 and can be introduced directly to the hydroformylation unit.

As for FIG. 1, the third hydrocarbon stream may exit the dehydrogenation-isomerization unit and enter the hydroformylation unit 120 through the third conduit 118. The hydroformylation unit 120 may have several entry points to allow for the input of additional process streams. Used in the description, the expression "flow entering the hydroformylation unit" is defined as the flow of process flows to the hydroformylation unit through one or more entry points. Examples of such process streams include, but are not limited to, additional streams from a dehydrogenation-isomerization unit 114, a hydrocarbon diluent stream, gas and / or other hydrocarbon streams, which may include olefins and paraffins obtained in other processes.

In the process of hydroformylation, olefins are converted to aldehydes, alcohols, or mixtures thereof due to the interaction of at least part of the olefins with carbon monoxide and hydrogen in accordance with the oxo process. As used herein, the term “oxo process” refers to the reaction of an olefin with carbon monoxide and hydrogen in the presence of a metal catalyst (eg, a cobalt catalyst) to produce an alcohol containing one carbon atom more than the starting olefin. Other hydroformylation processes use a “modified oxo process”. As used herein, “modified oxo process” refers to an oxo process in which a cobalt or rhodium catalyst modified with a phosphine, phosphite, arsine or pyridine ligand is used. The preparation and use of a modified oxo catalyst is described in US Pat. No. 3,231,621 (Slaugh), entitled "Reaction Rates in Catalytic Hydroformylation", in US Pat. No. 3,239,566 (Slaugh et al.) Under the name "Hydroformylation of Olefins," in US Pat. No. 3,239,569 (Slaugh et al.) Called “Hydroformylation of Olefins” in US Pat. No. 3,239,570 (Slaugh et al.) Called “Hydroformylation of Olefins” in US Pat. No. 3,239,571 (Slaugh et al.) Called “Hydroformylation of Olefins ", In US patent No. 3400163 (Mason et al.) Called" Bicyclic Heterocyclic Sec- and tert-Phoshines ", in US patent No. 3420898 (Van Winkle et al.) Called" Single Stage Hydroformylation of Olefins to Alcohols Single Stage Hydroformyla tion of Olefins to Alcohols ”, in US patent No. 3440291 (Van Winkle et al.) called“ Single Stage Hydroformylation of Olefins to Alcohols ”, in US patent No. 3448157 (Slaugh et al.) called“ Hydroformylation of Olefins ”, in US patent No. 3488158 (Slaugh et al.) under the name “Hydroformylation of Olefins”, in US patent No. 3496203 (Morris et al.) under the name “Tertiary Organophosphine-Cobalt-Carbonyl Complexes”; US Pat. No. 3,496,204 (Morris et al.) entitled “Tertiary Organophosphine-Cobalt-Carbonyl Complexes”; US Pat. No. 3,501,515 (Van Winkle et al.) entitled “Bicyclic Heterocyclic Terteriary Phoshine-Cobalt-Carbonyl Complexes”; US Pat. No. 3,527,818 (Mason et al.) entitled “Oxo Alcohols Using Catalysts Comprising Ditertiary Phosphines”; in US patent application, registration No. 10/075682 entitled “A Process for Preparing a Branched Olefin, a Method of Using the Branched Olefin for Making a Surfactant and a Surfactant” and in US patent application, registration No. 10/167209 called “ Process for Preparing of a Highly Linear Alcohol Composition. " Methods for producing alcohol are also described by Othmer in the publication “Encyclopedia of Chemical Technology” 2000, Fourth Edition; and Wickson in the publication “Monohydric Alcohols; Manufacture, Applications and Chemistry ”, Ed. Am. Chem. Soc. , 1981.

The hydroformylation catalyst used in hydroformylation unit 130 may include a metal from Group VIII of the Periodic Table. Examples of Group VIII metals are cobalt, rhodium, nickel, palladium or platinum. Group VIII metal can be used as a complex compound. The complex compound may be a Group VIII metal combined with a ligand. Examples of ligands are, but are not limited to, a phosphine, phosphite, arsine, stibine or pyridine ligand. Examples of hydroformylation catalysts include, but are not limited to, a cobalt-hydrocarbonyl catalyst, a cobalt-phosphine ligand catalyst, a rhodium-phosphine ligand catalyst, or combinations thereof.

In a hydroformylation unit 120, olefins in a third hydrocarbon stream may be hydroformylated using a continuous, batch, or batch process. In the case of a continuous operation, the hourly volumetric flow rates of the liquid can be from 0.1 to 10 hour ^-1 . When the hydroformylation unit 120 is operating as a batch process, the reaction time may vary from 0.1 to 10 hours or even more.

The temperatures in the hydroformylation unit 120 are from 100 to 300 ° C. In some embodiments of the invention, temperatures in a hydroformylation unit of 125 to 250 ° C. may be used. The pressure in the hydroformylation unit 120 may be from 1 to 300 atmospheres (from 101 to 30398 kPa). In one embodiment of the invention, a pressure of from 20 to 150 atmospheres (from 2027 to 15199 kPa) can be used. The amount of catalyst relative to the amount of olefin in the third hydrocarbon stream may be from 1: 1000 to 10: 1. In some embodiments, a ratio of from 1:10 to 5: 1 may be used. In one embodiment of the invention, a diluent stream may be added to hydroformylation unit 120 to control reaction conditions. The diluent stream may include solvents that do not substantially interfere with the target reaction. Examples of such solvents include, but are not limited to, alcohols, ethers, acetonitrile, sulfolane and paraffins.

At a hydroformylation unit 120, selectivity for monohydric alcohols of at least 90% and even at least 92% can be achieved. In addition, the degree of conversion of olefins to aliphatic alcohols can be from 50 wt.% To more than 95 wt.%. In some embodiments of the invention, the degree of conversion of olefins to aliphatic alcohols can be more than 75% of the mass. In some embodiments of the invention, the degree of conversion of olefins to aliphatic alcohols may be more than 99 wt.%.

Isolation of aliphatic alcohols obtained by the hydroformylation reaction of a product stream can be achieved by conventional known methods. In one embodiment of the invention, the isolation of aliphatic alcohols includes the first distillation of the obtained aliphatic alcohols, saponification, washing with water and re-distillation.

The hydroformylation reaction mixture stream may enter a separator 130 through a fourth conduit 132. In the separator 130, the hydroformylation reaction product stream may be subjected to a first distillation (for example, a single rapid distillation or molecular distillation). In one embodiment, molecular distillation can be used to produce at least two streams, a bottom stream and a overhead stream. In some embodiments, at least a portion of the bottom stream may be recycled to the hydroformylation unit 120 via a bottom stream recycling line 134. The overhead stream may contain, but is not limited to, paraffins, unreacted olefins, and a crude product containing aliphatic alcohols.

In one embodiment, the overhead stream may be saponified to remove any acids and esters present. Saponification can be carried out by reacting the head stream with an aqueous solution of a hydroxide base (for example, sodium hydroxide or potassium hydroxide) at elevated temperatures with stirring. Saponification can be carried out by the interaction of the head stream with a 0.5-10% aqueous hydroxide solution at a ratio of crude alcohol / water from 10: 1 to 1: 1. The amount of hydroxide used may depend on the estimated amount of esters and acids present.

Saponification of the head stream can be carried out periodically or continuously. The head stream may be subjected to one or more saponification processes. Saponification reaction temperatures can range from 40 to 99 ° C. In one embodiment of the invention, the saponification temperature may be from 60 to 95 ° C. Mixing the head stream with a layer of main water can be carried out during the saponification process. Separation of the head stream from the base water layer can be carried out using known methods. The head stream after separation may be flushed with water to remove any sodium salts present. The overhead stream can be separated using conventional known techniques (e.g., fractional distillation) to produce at least two streams, a crude alcohol product stream and a paraffin stream and unreacted olefins. Used in the description, the definition of "fractional distillation" refers to the distillation of liquids and the subsequent collection of fractions, determined by the boiling point. The stream of paraffins and unreacted olefins can be recycled, delivered to other processing plants, stored at the site, transferred to another facility and / or sold.

In some embodiments of the invention, the crude aliphatic alcohol product stream may contain undesired by-products (e.g., aldehydes, hemi-acetals). By-products can be removed due to the fact that the crude alcohol product stream undergoes a hydrotreatment step to produce an aliphatic alcohol stream. Used in this case, the definition of "hydrotreating" refers to the hydrogenation reaction carried out under relatively mild conditions. Hydrotreating can be carried out using conventional hydrogenation processes. Conventional hydrogenation processes may comprise passing the crude alcohol feed together with a stream of hydrogen over a bed of a suitable catalyst. The product stream may contain more than 50 wt.% Obtained aliphatic alcohols. In some embodiments of the invention, the product stream may contain more than 80 wt.% Obtained aliphatic alcohols. In some embodiments, the product stream may contain more than 95% by weight of aliphatic alcohols obtained. The product stream may contain branched aliphatic primary alcohols. The obtained aliphatic alcohols in the product stream can be sold, transferred to another company, stored on site and / or sent for use in other technological installations through the product pipeline 136.

The composition of the aliphatic alcohol product stream may contain hydrocarbons with an average carbon number of 8 to 19. In some embodiments, the average number of carbon atoms in the aliphatic alcohol product stream may be from 10 to 17. In some embodiments, the average number of carbon atoms of the feed stream can be from 10 to 13. In other embodiments, the average number of carbon atoms of the feed stream may be from 14 to 17. In some embodiments, branched aliphatic alcohol products may be used as a precursor for the production of anionic sulfates, including aliphatic sulfates and hydroxyalkyl sulfates and hydroxyalkyl alcohols.

Aliphatic alcohols may have a slightly higher aliphatic branching and a slightly higher number of quaternary carbon atoms as an olefin precursor. In some embodiments, the aliphatic branch may include methyl and / or ethyl branches. In other embodiments, the aliphatic branch may contain methyl, ethyl, and higher molecular weight aliphatic branches. In some embodiments of the invention, the number of quaternary carbon atoms in the aliphatic alcohol product may be less than 0.5. In other embodiments, the number of quaternary carbon atoms in the aliphatic alcohol product may be less than 0.3. Branches of the alcohol product can be determined using ¹ H NMR spectroscopy. The number of quaternary carbon atoms can be determined using ¹³ C NMR spectroscopy. The ¹³ C NMR method for determining quaternary carbon atoms for branched aliphatic alcohols is described in US Pat. No. 6,150,322 (Singleton et al.) Entitled “Highly Branched Primary Alcohol Compositions and Biodegradable Detergents Made Therefrom”.

In some embodiments of the invention, at least a portion of the paraffin and unreacted olefin stream may be combined with a second hydrocarbon stream in second conduit 116 to produce a combined stream through fifth conduit 138. The combined stream may be introduced into dehydrogenation-isomerization unit 114 through second conduit 116 At least a portion of the olefins in the combined stream may be isomerized to branched olefins. In some embodiments of the invention, at least a portion of the paraffin and unreacted olefin stream is introduced directly to the dehydrogenation-isomerization unit 114 through one or more entry points. Since a stream of paraffins and unreacted olefins containing paraffins and unreacted olefins can be recycled to the dehydrogenation-isomerization unit 114 as a single stream, the process can be more efficient, providing a higher total productivity. Higher performance will increase the overall yield of aliphatic alcohols.

In some embodiments of the invention, at least a portion of the paraffins and unreacted olefins stream may be introduced to the hydrogenation unit 110 via a fifth conduit 138. In some embodiments of the invention, at least a portion of the paraffins and unreacted olefins stream may be combined with the first hydrocarbon stream in line 112 through the fifth line 138. The combined stream may enter the hydrogenation unit 110 and continue the process of producing aliphatic alcohols.

In some embodiments, the concentration of olefins and paraffins in hydroformylation unit 120 may be adjusted depending on the source of olefin stream entering the hydroformylation unit. A fourth hydrocarbon stream may be added before entering the hydroformylation unit 120 to form a combined stream. In another embodiment, a fourth hydrocarbon stream may be introduced directly to hydroformylation unit 120 through one or more points. A fourth hydrocarbon stream may be introduced into the third pipe 118 through the sixth pipe 140 to form a combined stream. The combined stream may enter hydroformylation unit 120 through a third conduit 118 to continue the process of producing aliphatic alcohols.

The fourth hydrocarbon stream may come from the same source as the first hydrocarbon stream. In some embodiments, the fourth hydrocarbon stream may be a stream containing olefins, paraffins and / or hydrocarbon solvents obtained from another source. The fourth hydrocarbon stream may contain olefins and paraffins. In some embodiments of the invention, the average number of carbon atoms in the fourth hydrocarbon stream is from 7 to 18. In some embodiments, the paraffin content in the fourth hydrocarbon stream may be from 60 to 90% by weight. In other embodiments of the invention, the paraffin content in the fourth hydrocarbon stream may be more than 90 wt.%.

In one embodiment of the invention, the olefin content in the fourth hydrocarbon stream is from 1 to 99% relative to the total hydrocarbon content. In some embodiments, the olefin content in the fourth hydrocarbon stream may be from 45 to 99 wt.%. In other embodiments of the invention, the concentration of olefins in the fourth hydrocarbon stream may be more than 80 wt.%.

In some embodiments of the invention, the dehydrogenation-isomerization unit 114 can be divided into a large number of zones in order to control reaction temperatures and / or prevent unwanted side reactions (for example, the formation of dienes and / or cyclization reactions). The first hydrocarbon stream containing paraffins and unreacted olefins can be introduced into the hydrogenation unit 110 through the first pipe 112, as shown in the case of the system 200 of FIG. 3A. Hydrogenation of olefins can proceed as in the previously described hydrogenation unit 110 to produce a second hydrocarbon stream. The second hydrocarbon stream may exit hydrogenation unit 110 through second conduit 116 and enter the dehydrogenation-isomerization unit 114. The dehydrogenation-isomerization unit 114 can be divided into a large number of zones. A large number of zones may include, but are not limited to, a first reaction zone, a transition zone, and a second reaction zone. In the first reaction zone 210, at least a portion of the paraffins in the second hydrocarbon stream may be dehydrogenated to olefins. The process stream may then flow to the second reaction zone 212. In the second reaction zone 212, at least a portion of the olefins in the process stream can be isomerized to branched olefins to produce a third hydrocarbon stream.

In the first reaction zone 210, the dehydrogenation catalyst may be selected from a wide variety of types of catalysts. For example, the catalyst may be based on a metal or a metal compound supported on a porous substrate. The metal or metal compound may be selected, but not limited to, from chromium oxide, iron oxide and noble metals.

Methods of preparing catalysts for carrying out the dehydrogenation step and for carrying out the separation steps associated with them are known in the art. For example, suitable methods for preparing catalysts and carrying out the dehydrogenation step are described in US Pat. No. 5,012,021 (Vora et al.) Entitled “Process for Production of AlkylAromatic Hydrocarbons Using Solid Catalysts”, in US Pat. No. 3,274,287 (Moore et al.) Called “ Hydrocarbon Conversion Process and Catalyst ”, in US patent No. 3315007 (Abell et al.) Called“ Dehydrogenetion of Saturated Hydrocarbons Over Noble-Metal Catalyst ”, in US patent No. 3315008 (Abell et al.) Under the name“ Dehydrogenetion of Saturated Hydrocarbons Over Noble-Metal Catalyst ”, in US Pat. No. 3,745,112 (Rausch), entitled“ Platinum-Tin Uniformly Dispersed Hydrocarbon Conversion Catalyst and Process ”, in US Pat. No. 4,506,032 (Imai et al.) Under the name iem "Dehydrogenation Catalyst Composition" and U.S. Patent 4430517 № (Imai et al.), entitled "Dehydrogenation Process Using a Catalytic Composition".

Temperatures in the first reaction zone 210 may range from 300 to 600 ° C. In some embodiments of the invention, the reaction temperature in the first reaction zone 210 may be from 450 to 550 ° C. The total pressure in the first reaction zone 210 may be from 0.010 to 25.0 atmospheres (from 1 to 2534 kPa). In some embodiments, the total pressure in the first reaction zone 210 may be from 0.010 to 15.0 atmospheres (from 1 to 1520 kPa). In some embodiments of the invention, hydrogen may be supplied together with the unreacted second hydrocarbon stream to prevent coking. Hydrogen and paraffins present in the unreacted second hydrocarbon stream may be supplied at a molar ratio of hydrogen to paraffin from 0.1 to 20. In one embodiment of the invention, the molar ratio of hydrogen to paraffin may be from 1 to 10.

The residence time in the first reaction zone 210 can be selected so that the degree of conversion of paraffins to olefins is below 50 mol%. In some embodiments of the invention, the degree of conversion of paraffins to olefins can be maintained from 10 to 20 mol%. By maintaining low conversion values, side reactions (e.g. diene formation and cyclization reactions) can be prevented. In some embodiments of the invention, the olefin hydrocarbon stream may exit the first reaction zone 210, pass through the transition zone 214, and enter the second reaction zone 212. The transition zone 214 may comprise a heat exchanger 216. The heat exchanger 216 may lower the temperature of the olefinic hydrocarbon stream. In one embodiment of the invention, the first reaction zone 210 and the second reaction zone 212 of the dehydrogenation-isomerization unit 114 may be separate units, as shown in FIG. 3B, with a heat exchanger 216 located between them.

After the olefin hydrocarbon stream enters the second reaction zone 212, at least a portion of the olefins isomerized to branched olefins to produce a third hydrocarbon stream. The composition and degree of branching of the third hydrocarbon stream can be determined using ¹ H NMR spectroscopy. In one embodiment, the olefin stream may exit the first reaction zone 210 and directly enter the second reaction zone 212, where at least a portion of the olefins in the olefin stream is isomerized to branched olefins.

The catalyst used to isomerize olefins to branched olefins may be the same catalyst as described in US Pat. No. 5,648,584 (Murray) under the name “Process for Isomerizing Linear Olefins to Isoolefins” and US Pat. No. 5,648,585 (Murray et al.) titled “Process for Isomerizing Linear Olefins to Isoolefins”.

In an embodiment of the invention, linear olefins in the second hydrocarbon stream are isomerized in the second reaction zone 212 by reacting at least a portion of the olefin stream with a zeolite catalyst. The zeolite catalyst may have at least one channel with a crystallographically free diameter from more than 4.2 Å to less than 7 Å. The zeolite catalyst may have a sufficiently large elliptical pore size to allow the linear olefin to enter and at least partial diffusion of the branched olefin. The pore size of the zeolite catalyst may also be small enough to slow the formation of coke.

The temperature in the second reaction zone 212 can range from 200 to 500 ° C. in order to linearize the olefins to branched olefins. In some embodiments of the invention, the reaction temperatures in the first reaction zone and in the second reaction zone are essentially the same. In such embodiments, the use of a heat exchanger is not required. However, typically the reaction temperature of the second reaction zone 212 is lower than the reaction temperature of the first reaction zone. The use of a heat exchanger reduces the temperature of the stream leaving the first reaction zone to a temperature suitable for the reaction in the second reaction zone. The partial pressure of the hydrocarbon in the second reaction zone may be from 0.1 to 10 atmospheres (from 10 to 1013 kPa).

In some embodiments, the third hydrocarbon stream may exit the second reaction zone and enter a separation unit. In a separation unit, branched olefins can be separated from linear olefins and paraffins, as shown previously in FIG. 1. With reference to FIG. 3A, the third hydrocarbon stream may exit the second reaction zone 212 through the third pipe 118 and enter a hydroformylation unit 120. At least a portion of the olefins in the third hydrocarbon stream may be hydroformylated to produce a hydroformylation reaction stream, such as described for system 100 of FIG. 1. At least a portion of the hydroformylation reaction stream may be separated into a lower stream and a overhead stream using conventional known methods. The crude aliphatic alcohol product stream may be further purified as described for system 100 to produce a paraffin and unreacted olefin stream and an aliphatic alcohol product stream. The aliphatic alcohol product stream may contain branched aliphatic alcohols (e.g., branched primary aliphatic alcohols). The product stream of aliphatic alcohol can be recycled, transferred to other processing plants, sold and / or transferred to storage tanks.

The hydroformylation reaction mixture stream may enter separator 130 through a fourth conduit 132. In separator 130, at least three streams, a bottom stream, a paraffin and unreacted olefin stream and an aliphatic alcohol product stream can be obtained using the procedures previously described for the system 100. At least a portion of the bottom stream may be recycled to the hydroformylation unit through a bottom stream recycling pipeline 134. At least a portion of the paraffin and unreacted olefin stream may t be recycled, combined with other process streams, sent to other plant and / or directed to on-site storage. The aliphatic alcohol product stream may be transferred via line 136 for on-site storage, industrial sale, transfer to another facility, and / or for use in other process plants.

In one embodiment, at least a portion of the paraffin and unreacted olefin stream may be combined with a second hydrocarbon stream to produce a combined hydrocarbon stream through a fifth conduit 138. The combined hydrocarbon stream may enter the first reaction zone 210 and undergo a dehydrogenation-isomerization process and a process hydroformylation to obtain aliphatic alcohols. By recirculating the flow of paraffins and unreacted olefins, the yield of the product can be maximized. In one embodiment of the invention, a stream of paraffins and unreacted olefins can be directly introduced to the dehydrogenation-isomerization unit 114 through one or more entry points.

In some embodiments of the invention, the concentration of olefins and paraffins in hydroformylation unit 120 may be used depending on the source of olefin stream entering the hydroformylation unit, as previously described for system 100. At least a portion of the fourth hydroformylation stream may be introduced into the third conduit 118 the flow of the hydroformylation unit 120 through the sixth conduit 140. The combined flow may be introduced into the hydroformylation unit 120 through the third conduit d 118. At least a portion of the olefins in the combined stream may be hydroformylated to produce aliphatic alcohols. In one embodiment, a third hydrocarbon stream may be introduced directly to hydroformylation unit 120 through one or more entry points.

A fourth hydrocarbon stream may be used to optimize the concentration of olefins in hydroformylation unit 120 at a concentration sufficient to maximize hydroformylation of the olefin. In addition, the fourth hydrocarbon stream can optimize the ratio of linear groups to branched aliphatic groups in the aliphatic alcohol. The fourth hydrocarbon stream may be, but is not limited to, a hydrocarbon stream containing olefins, paraffins and / or hydrocarbon solvents.

In one embodiment, the catalyst in the dehydrogenation-isomerization unit 114 can be used in the configuration of the composite layer. The configuration of the composite layer allows the use of one or more catalysts in the reactor. A paraffin dehydrogenation catalyst and an olefin isomerization catalyst can increase the selectivity of the catalyst system and / or process. The configuration of the composite layer of the dehydrogenation-isomerization unit 114 is shown for the system 300 of FIG. 3. The operating conditions for the configuration of the composite layer may be the same as for the two-zone system described in the case of the system 200. The second hydrocarbon stream may enter the dehydrogenation zone 310 through the second pipe 116.

The dehydrogenation catalyst used in the configuration of the composite layer may have non-acidic properties. The definition of “non-acidic properties” used in this case refers to a catalyst that exhibits little skeletal isomerization activity. The dehydrogenation catalyst may comprise a noble metal, a group IVA component, an alkali or alkaline earth metal component, a halogen component and / or a porous carrier material.

In some embodiments of the invention, the noble metal can be distributed throughout the dehydrogenation catalyst. The amount of the noble metal can be from 0.01 to 5 wt.% Calculated on the element and calculated on the final dehydrogenation catalyst. In some embodiments of the invention, the dehydrogenation catalyst contains from 0.1 to 1 wt.% Platinum. The noble metal can be introduced into the catalytic system using methods known in the art (for example, co-precipitation, co-gelation, ion exchange, impregnation, vapor deposition or from a source of atoms) before introducing other catalytic components. In some embodiments of the invention, the noble metal may be introduced into the catalyst system while impregnating with other catalyst components. In other embodiments of the invention, the noble metal may be introduced into the catalyst system after the introduction of other catalyst components. In some embodiments of the invention, the noble metal can be introduced by impregnating the support material with a solution or suspension of a decomposable noble metal compound. For example, platinum can be added to the catalyst support by extracting platinum from an aqueous solution of platinum chloride. In another embodiment, optional components (eg, nitric acid) may be added to the impregnation solution to facilitate the distribution or fixation of the noble metal in the final catalyst component.

The component of group IVA may be germanium, tin, lead, or a combination thereof. In some embodiments of the invention, a group IVA component may exist within the catalyst in a more oxidized state than the noble metal. A component of group IVA may be present as an oxide. In some embodiments, an IVA component may be combined with a carrier material. In some embodiments of the invention, the component of the group IVA can be combined with other catalytic components. In other embodiments, an IVA component may be distributed throughout the catalyst. The component of group IVA can be from 0.01 to 5 wt.% Calculated on the element and the final catalytic system. In some embodiments, the catalyst contains from 0.2 to 2 wt.% Tin.

The component of group IVA can be introduced into the catalyst system in accordance with conventional known methods (e.g., co-precipitation, co-gelation, ion exchange and impregnation) prior to the introduction of other catalytic components. In some embodiments of the invention, the component of the group IVA can be introduced together with other catalytic components. In other embodiments, an IVA component may be administered after administration of other catalytic components. In some embodiments, the tin component may be introduced via co-gelation with a porous carrier material. Tin can be combined with an alumina carrier material by mixing a soluble tin compound (e.g., tin dichloride or tin tetrachloride) with an aluminum oxide hydrosol. A gelling agent (e.g., hexamethylene-tetraamine) may be added to the tin (alumina hydrosol) mixture. The tin- (alumina hydrosol) mixture can be added dropwise to an oil bath to form spheres containing alumina and tin. In some embodiments of the invention, the germanium component can be impregnated with a carrier material using a solution or suspension of a decomposable germanium compound (for example, germanium tetrachloride dissolved in alcohol). In other embodiments, a lead component may be applied by impregnation from a solution of lead nitrate in water.

In some embodiments of the invention, an alkali or alkaline earth metal component may be introduced into the dehydrogenation catalyst. The alkali or alkaline earth metal component may be, but is not limited to, cesium, rubidium, potassium, sodium, lithium, barium, strontium, calcium and magnesium, or mixtures thereof. A component of an alkali or alkaline earth metal may be present in the final catalytic composite in an oxidation state higher than the noble metal. The alkali or alkaline earth metal component may be present as an oxide. In some embodiments of the invention, the alkali or alkaline earth metal may be combined with a carrier material. In some embodiments, the alkali or alkaline earth metal may be combined with other catalytic dehydrogenation components.

In other embodiments of the invention, the alkaline or alkaline earth metal component may be distributed throughout the catalyst system. The amount of alkaline or alkaline earth component may be from 0.01 to 15 wt.% Based on the element and the final catalytic composite. In other embodiments of the invention, the dehydrogenation catalyst contains from 1 to 3 wt.% Potassium. In some embodiments, the atomic ratio of the alkaline or alkaline earth component in the noble metal is at least greater than 10.

The alkaline or alkaline earth component can be introduced into the catalytic system by generally known methods (for example, co-precipitation, co-gelation, ion exchange and impregnation) before the introduction of other catalytic components. In some embodiments of the invention, the alkali or alkaline earth metal component may be introduced together with other catalytic components. In other embodiments, the alkaline or alkaline earth component may be added after the introduction of other catalytic components. For example, potassium can be introduced by impregnating the support material with a potassium nitrate solution. The atomic ratio of the alkaline or alkaline earth component to the noble metal may be at least 10. In some embodiments, the atomic ratio of the alkaline or alkaline earth component to the noble metal can be from 15 to 25.

The porous carrier material used in the dehydrogenation catalyst may be a porous absorbent substrate with a high specific surface area of 25 to 500 m ² / g. The porous carrier material has a melting point greater than the temperature in the dehydrogenation zone. Examples of carrier materials include, but are not limited to, activated carbon, coke, carbon, silicon dioxide, silica gel, silicon carbide, synthetically prepared and / or naturally occurring clays and silicates, refractory inorganic oxides, crystalline zeolite aluminosilicates, existing in nature or synthetically prepared mordenite and / or faujasite, spinel, or combinations thereof. In some embodiments of the invention, the carrier material may be γ- or β-alumina. In some embodiments, clays and silicates may or may not be acid-treated (e.g., attapulgite, porcelain clay, diatomaceous earth, fuller's earth, kaolin, kieselguhr, ceramics, porcelain, crushed refractory bricks, bauxite). Examples of refractory inorganic oxides are alumina, titanium dioxide, zirconia, chromium oxide, beryllium oxide, vanadium oxide, cerium oxide, hafnium oxide, zinc oxide, magnesium oxide, boron oxide, thorium oxide, aluminosilicate, magnesium silicate, aluminosilicate and aluminosilicate . Zeolite aluminosilicates in some embodiments may be in hydrogen form. In other embodiments, zeolite aluminosilicates may be in a form that can be exchanged with metal cations. Examples of spinels are, but are not limited to, MgAl ₂ O ₄ , FeAl ₂ O ₄ , ZnAl ₂ O ₄ , CaAl ₂ O ₄ and other similar compounds having the formula MO-Al ₂ O ₃ in which M is a metal having a valency 2.

In some embodiments of the invention, the alumina carrier material used in the dehydrogenation catalyst may be prepared in an appropriate manner from synthetic or naturally occurring raw materials. The alumina carrier material can be obtained in any desired form (for example, in the form of spheres, balls, pellets, extrudates, powders, granules). The alumina carrier can be used in the form of any particles. In some embodiments, it can be used in the form of spheres. Particles may have a diameter of 1/16 inch. In some embodiments, a particle diameter of less than 1/32 inch may be used.

Alumina spheres can be prepared in some embodiments by converting aluminum metal to an alumina sol. An alumina sol can be obtained by reacting aluminum metal with suitable peptizing acid and water. The resulting alumina sol and gelling agent can be added dropwise to an oil bath to form spherical alumina gel particles. The resulting alumina gel can be converted into a carrier material based on γ- or β-alumina using known techniques (for example, by aging, drying and calcining).

In other embodiments of the invention, alumina cylinders can be prepared by mixing the alumina powder with water and a suitable peptizing agent (e.g., nitric acid) in a paste to form an extrudable composition. The composition can be extruded through a die of a suitable size, then cut into an extrudate. Other forms of alumina-based carrier material may be prepared by conventional methods. After forming the alumina particles, they can be dried and calcined. During the preparation, the alumina carrier can be subjected to intermediate treatments (for example, washing with water or with a solution of ammonium hydroxide).

The dehydrogenation catalyst may contain halogen. Halogen may be, but without limitation, fluorine, chlorine, bromine, iodine, or mixtures thereof. Halogen may be present in a combined state with the porous carrier material. In some embodiments of the invention, the halogen component may be distributed throughout the catalyst system. Halogen can be from 0.2 to 15 wt.% Based on the element and the final catalytic system. In some embodiments, the dehydrogenation catalyst contains from 1 to 3 wt.% Chlorine.

In some embodiments of the invention, the catalytic system may contain at least 0.2 wt.%, Calculated on the element, halogen. Halogen in the catalyst can improve activity in the case of dehydrogenation of hydrocarbons. In some embodiments, the active halogen component inhibits the formation of carbon on the catalyst during dehydrogenation. An advantage of the catalyst system may be that side reactions of undesired isomerization or coke formation can be slowed down. In some embodiments of the invention, halogen may increase the activity of the catalyst. Activity can be reduced by steam treatment of the dehydrogenation catalyst to remove excess halogen from the dehydrogenation catalyst.

Halogen can be introduced into the catalyst system by any suitable means. The introduction of halogen can be carried out before receiving the carrier material. In some embodiments, the introduction of halogen can be carried out during the introduction of other components. In other embodiments, the introduction of halogen can be carried out after the introduction of other catalytic components. In some embodiments of the invention, the alumina ash carrier comprises a halogen that can provide at least some of the halogen content of the final catalyst. In some embodiments of the invention, the halogen content or part of it can be added to the catalyst system while combining the carrier material with other catalyst components (for example, when using platinum chloride acid to impregnate the platinum component). In other embodiments of the invention, a halogen component or part thereof can be added to the catalyst system by contacting the halogen catalyst. In some embodiments, the halogen may be added to the catalyst in the form of a compound, solution, suspension or dispersion containing a halogen (eg, hydrochloric acid), before or after combining other catalyst components with a carrier material. In some embodiments, the halogen or part thereof may be introduced by contacting the catalyst with a compound, solution, suspension or dispersion containing a halogen in a subsequent catalyst regeneration step. In the regeneration step, carbon deposited on the catalyst in the form of coke during application of the catalyst in the hydrocarbon conversion step can be burned out of the catalyst. The noble metal on the catalyst can be redistributed again to obtain a regenerated catalyst with characteristics similar to those of fresh. Halogen can be added during the carbon extraction step or during the redistribution stage of the noble metal (for example, by contacting the catalyst with gaseous hydrogen chloride). In some embodiments, the halogen may be added to the catalyst system by adding a halogen or compound, solution, suspension or dispersion containing a halogen (e.g., propylene dichloride) to a hydrocarbon feed stream. In other embodiments of the invention, halogen may be added to the recycle gas during operation of the dehydrogenation unit.

In some embodiments of the invention, the dehydrogenation catalyst may contain sulfur from 0.01 to 10 wt.% Based on the element and the final catalyst composition. Sulfur may be introduced into the catalyst system in any suitable manner. In some embodiments of the invention, sulfur or a sulfur-containing compound (for example, hydrogen sulfide or a lower molecular weight mercaptan) can be contacted with the catalyst system in the presence of hydrogen at a temperature of from 10 to 540 ° C., under anhydrous conditions. The ratio of hydrogen to sulfur in some embodiments may be 100.

The dehydrogenation catalyst in some embodiments of the invention may contain other additional components or mixtures thereof, which act separately or in interaction, as modifiers, to improve the catalytic activity, selectivity or stability of the catalytic system. Examples of catalyst modifiers include, but are not limited to, antimony, arsenic, beryllium, bismuth, cadmium, calcium, chromium, cobalt, copper, gallium, gold, indium, iron, lithium, magnesium, molybdenum, nickel, rhenium, scandium, silver , tantalum, thallium, titanium, tungsten, uranium, zinc and zirconium. Catalyst modifiers may be added in any suitable manner to the carrier material during the preparation of the dehydrogenation catalyst. In other embodiments of the invention, the catalyst modifiers may be added in any suitable manner after the preparation of the dehydrogenation catalyst. In some embodiments of the invention, catalyst modifiers may be added in any suitable manner to the catalyst system prior to the introduction of other catalyst components. In some embodiments of the invention, catalyst modifiers may be added in conjunction with other catalytic components. In other embodiments, catalyst modifiers may be added after the introduction of other catalyst components. A description of the dehydrogenation catalyst can be found in US Pat. No. 4,506,032 (Imai et al.) Entitled “Dehydrogenation Catalyst Composition”.

The olefin hydrocarbon stream may enter the isomerization zone 312. In some embodiments, it may be necessary to lower the temperature from the dehydrogenation zone 310 to the isomerization zone 312 to prevent cracking of the olefin hydrocarbon stream when it enters the isomerization zone. Cold hydrogen gas can be introduced into the dehydrogenation zone 310 via a gas line 314 to control temperatures in the dehydrogenation zone 310. In the isomerization zone 312, at least a portion of the olefins in the olefinic hydrocarbon stream can be isomerized to branched olefins to produce a third hydrocarbon stream .

In some embodiments of the invention, the isomerization catalyst may be the same catalyst as described for the isomerization of olefins in system 200. A description of the isomerization catalyst can be found in US Pat. No. 5,510,306 (Murray), entitled “Process for Isomerizing Linear Olefins to Isoolefins”. In some embodiments of the invention, from 0.01 to 5% by weight of a noble metal may be added to the isomerization catalyst used in the configuration of the composite layer to increase the dehydrogenation activity of the zeolite catalyst. Conventional methods for introducing a metal (e.g., impregnation, ion exchange of a noble metal, co-kneading into a paste) to introduce a noble metal (e.g., platinum, palladium) into a zeolite can be used to obtain a working catalyst that can be used in the dehydrogenation-isomerization of paraffins.

A third hydrocarbon stream may exit the isomerization zone 312 and enter a hydroformylation unit 120 through a third conduit 118. At least a portion of the olefins in the third hydrocarbon stream may be hydroformylated to produce a reaction stream as described for system 100. At least a portion of the hydroformylation reaction stream may be divided into a lower stream and a overhead stream using generally known methods. The overhead stream can be cleaned and separated as described for system 100 to produce a paraffin stream and unreacted olefins and a crude aliphatic alcohol product stream. The crude aliphatic alcohol product stream may be further purified as described in the case of system 100 to obtain an aliphatic alcohol product stream. The hydroformylation reaction mixture stream may enter separator 130 through a fourth conduit 132. In separator 130, at least two streams, a bottom stream and a overhead stream, can be obtained as previously described for system 100. The bottom stream can be recycled to the hydroformylation unit 120 through a bottom stream recycling conduit 134. The overhead stream can be purified and separated into at least two streams, a paraffin and unreacted olefin stream and a crude aliphatic alcohol product stream. At least a portion of the paraffin and unreacted olefin stream may be recycled, combined with other process streams, sent to other process units, and / or sent to storage facilities. The crude aliphatic alcohol product stream may be further purified as described in the case of system 100 to obtain an aliphatic alcohol product stream. The product stream may contain branched aliphatic alcohols (for example, branched primary aliphatic alcohols). The aliphatic alcohol product stream can be transferred through product line 136 so that it can be stored on the spot, sold commercially, transferred to other plants and / or used in other process plants.

At least a portion of the paraffin and unreacted olefin stream may be combined with the second hydrocarbon stream in the second conduit 116 to produce a combined hydrocarbon stream through the fifth conduit 138. The combined hydrocarbon stream may enter the dehydrogenation zone 310 of the dehydrogenation-isomerization unit 114 through the second conduit 116. The combined hydrocarbon stream entering the dehydrogenation zone 310 enters the dehydrogenation-isomerization and hydroformylation processes to obtain an aliphatic their alcohols. By recirculating the flow of paraffins and unreacted olefins, the yield of the product can be maximized. In one embodiment of the invention, a stream of paraffins and unreacted olefins can be directly introduced to the dehydrogenation-isomerization unit 114 through one or more entry points.

In some embodiments, the concentration of olefins and paraffins in hydroformylation unit 120 may be controlled depending on the source of olefin stream entering the hydroformylation unit. At least a portion of the fourth hydrocarbon stream may be introduced into the third conduit 118 before entering the hydroformylation unit, as previously described for system 100. The combined stream may be introduced into the hydroformylation unit 120 via conduit 118. At least a portion of the olefins in the combined the stream may be hydroformylated to produce aliphatic alcohols.

Aliphatic alcohols produced in systems 100, 200, and 300 can be converted to oxo alcohols, sulfates, or other industrial products. At least a portion of the aliphatic alcohols in the aliphatic product stream can be reacted in an oxyalkylation unit using epoxides (e.g. ethylene oxide, propylene oxide, butylene oxide) in the presence of a base to produce hydroxyalkyl alcohol. The condensation of alcohol with epoxides allows the alcohol functionality to increase in size due to one or more hydroxy groups. The number of oxy groups can range from 3 to 12. For example, the reaction of an alcohol with ethylene oxide can produce alcohol products having from 3 to 12 ethoxy groups. The reaction of an alcohol with ethylene oxide and propylene oxide can give alcohols with an ethoxy / propoxy ratio of ethoxy groups to propoxy groups from 4: 1 to 12: 1. In some embodiments of the invention, a significant portion of the alcohol moieties is combined with more than three ethylene oxide moieties. In other embodiments, approximately equal parts may be mixed with less than three ethylene oxide moieties. In a typical mixture of oxyalkylation products, a small portion of unreacted alcohol may be present. In an embodiment of the invention, at least a portion of the aliphatic alcohol product stream can be obtained by condensation of a C ₅ -C ₃₁ aliphatic alcohol with an epoxide. In some embodiments, branched primary C ₅ -C ₁₅ alcohols may be condensed with ethylene oxide and / or propylene oxide. In other embodiments, branched primary C ₁₁ -C ₁₇ alcohols may be condensed with ethylene oxide and / or propylene oxide. The resulting oxyalkyl alcohols can be sold commercially, transferred to other enterprises, stored on site and / or used in other process plants. In some embodiments, the hydroxyalkyl alcohol may be sulfonated to give an anionic surfactant.

In one embodiment of the invention, at least a portion of the alcohols in the aliphatic alcohol product stream may be added to the base. The base may be an alkali metal or alkaline earth metal hydroxide (e.g., sodium hydroxide or potassium hydroxide). The base may act as a catalyst for the oxyalkylation reaction. For the hydroxyalkylation of alcohol, an amount of 0.1 to 0.6% by weight of base can be used, based on the total weight of the alcohol. In one embodiment of the invention, the mass percentage of the base may be from 0.1 to 0.4 wt.% Based on the total amount of alcohol. The reaction of the alcohol with a base can produce alkoxide. The resulting alkoxide can be dried to remove any water present. Dried alkoxide can be reacted with epoxide. The amount of epoxide used can be from 1 to 12 moles per mole of alkoxide. The resulting mixture (alkoxide-epoxide) should react until the epoxide is used up. A decrease in overall reaction pressure may indicate that the reaction has ended.

The temperatures in the oxyalkylation unit are from 120 to 220 ° C. In one embodiment, the reaction temperature may be from 140 to 160 ° C. The reaction pressure can be achieved by introducing the required amount of epoxide into the reaction vessel. Epoxides have a high vapor pressure at the desired reaction temperature. From the point of view of process safety, the partial pressure of epoxy reagents should be limited, for example, to less than 4 atmospheres (413 kPa). Other safety measures may include diluting the reagents with an inert gas such as nitrogen. For example, inert gas dilution may result in a vapor phase concentration of reactants of 50% or less. In some embodiments of the invention, the alcohol-epoxide reaction can be safely carried out at a higher concentration of the reagent, a higher total partial pressure and a higher partial pressure of the epoxide, if suitable, generally known safety measures are taken to control the explosion safety. As for ethylene oxide, a total pressure of 3 to 7 atmospheres (304 to 709 kPa) can be used. The total pressure of ethylene oxide from 1 to 4 atmospheres (from 101 to 415 kPa) can be used in some embodiments of the invention. In one embodiment of the invention, a total pressure of 1.5 to 3 atmospheres (150 to 304 kPa) relative to ethylene oxide may be used. Pressure is a measure of the degree of reaction. The reaction can be considered essentially complete when the pressure does not drop in time.

Aliphatic alcohols and oxyalkyl alcohols can be derivatized to produce compositions (e.g., sulfonates, sulfates, phosphates) that can be used in formulations of commercial products (e.g., detergents, surfactants, oil additives, lubricating oil formulations). For example, alcohols can be sulfonated with SO ₃ to produce sulfates. The term “sulfonated” refers to a sulfur atom or a sulfur-containing functionality added to a carbon or oxygen atom. Sulfonation processes are described in US Pat. No. 6,462,215 (Jacobson et al.) Under the name “Sulfonation, Sulfation and Sulfamation”, in US Pat. No. 6,448,435 (Jacobson et al.) Under the name “Sulfonation, Sulfation and Sulfamation”, in US Pat. No. 3462525 (Levinsky et al.) Under the name “Dental Compositions Comprising Long-Chain Olefin Sulfonates” in US Pat. No. 3,428,654 (Rubinfeld et al.) Under the name “Alkene Sulfonation Process and Products” in US Pat. No. 3,420,875 (DiSalvo et al. ) under the name “Olefin Sulfonates”, in US patent No. 3506580 (Rubinfeld et al.) under the name “Heat-Treatment of Sulfonated Olefin Products” and in US patent No. 3579537 (Rubinfeld) under the name “Process for Separation of Sulfones from Alkenyl Sulfonic Acids. "

The general class of sulfates of aliphatic alcohols can be characterized by the chemical formula: (RO- (A) _x —SO ₃ ) _n M. R is an aliphatic residue. “A” represents an alkylene oxide residue; x corresponds to the average number of residues "A" per residue R-O and may be from 0 to 15; and n is a number depending on the valency of cation M. Examples of cation M are, but are not limited to, alkali metal ions, alkaline earth metal ions, ammonium ions and / or mixtures thereof. Examples of cations are, but are not limited to, magnesium, potassium, monoethanolamine, diethanolamine, or triethanolamine.

Aliphatic and hydroxyalkyl alcohols can be sulfated at the facility. The sulfation process involves the reaction of sulfur trioxide (SO ₃ ), chlorosulfonic acid (ClSO ₃ H), sulfamic acid (NH ₂ SO ₃ H) or sulfuric acid with alcohol. In one embodiment of the invention, sulfur trioxide in concentrated (e.g., fuming) sulfuric acid can be used to sulfate alcohols. Concentrated sulfuric acid may have a concentration of from 75 to 100 wt.%. In one embodiment, sulfuric acid may have a concentration of from 85 to 98 wt.%. The amount of sulfur trioxide may be from 0.3 to 1.3 moles of sulfur trioxide per mole of alcohol. In some embodiments of the invention, sulfur trioxide may be from 0.4 to 1.0 moles of sulfur trioxide per mole of alcohol.

In one embodiment of the invention, a sulfur trioxide sulfation technique may include reacting liquid alcohol or hydroxyalkyl alcohol and sulfur dioxide gas in a sulfator with a flowing film to produce a sulfuric acid ester and an alcohol. The reaction zone of the sulphation unit with a falling film can operate at atmospheric pressure and at a temperature of 25 to 70 ° C. The sulfuric acid and alcohol ester may exit the sulfation unit with a falling film and enter the neutralization reactor. The sulfate ester can be neutralized with an alkali metal solution to give an alkyl sulfate salt or an oxyalkyl sulfate salt. Examples of alkali metal solutions are sodium or potassium hydroxide solutions.

Derivative alcohols can be widely used. Examples of such applications are detergents. Detergents are, but are not limited to, granular and liquid detergents for washing clothes, liquid detergents for washing dishes and other preparations. Examples of other drugs include general purpose cleansing agents, liquid soaps, shampoos, and liquid cleaners.

Granular laundry detergents may include a number of components besides derivatized alcohols (e.g. surfactants, modifiers, concomitant modifiers, bleach agents, bleach activators, foam control agents, enzymes, anti-bleaching agents, optical non-bleaching agents and stabilizers). Examples of surfactants can be ionic, nonionic, amphoteric or cationic surfactants.

Liquid detergents for washing clothes include the same components as granular detergents for washing clothes. In some embodiments of the invention, liquid laundry detergents may include less of an inorganic modifying agent component than granular laundry detergents. Hydrotropic compounds may be present in liquid detergents. General-purpose cleaning agents may include other surfactants, modifying agents, foam regulators, hydrotropes, and solubilizing alcohols.

EXAMPLES

Example 1: Isomerization of olefins in a hydrocarbon stream obtained in the Fischer-Tropsch process

Carbon monoxide and hydrogen are reacted under the conditions of the Fischer-Tropsch process to produce a hydrocarbon mixture of linear paraffins, linear olefins, a small amount of dienes and a small amount of oxygenates. The Fischer-Tropsch hydrocarbon stream is separated into various hydrocarbon streams using fractional distillation. A hydrocarbon stream is obtained containing olefins and paraffins with an average number of carbon atoms from 8 to 10. The composition of the obtained C ₈ -C ₁₀ hydrocarbon stream is analyzed by gas chromatography, the results are shown in table 1.

Table 1 The composition of the hydrocarbon stream of the Fischer-Tropsch process Wt% C ₇ and lighter hydrocarbons 0.12 ₈ branched olefins 0.02 With ₈ linear olefins 0.75 1 octen 0.69 n-octane 2.21 ₉ branched olefins 0.16 With ₉ linear olefins 8.52 1-Nonen 8.07 n-nonan 20.03 With ₁₀ branched olefins 0.28 With ₁₀ linear olefins 22.92 1-decen 20.87 n-dean 41.12 C ₁₁ and heavier hydrocarbons 0.21 C ₉ -C ₁₁ alcohols 3.56

The zeolite catalyst used for the isomerization of linear olefins in a hydrocarbon stream is prepared as follows. Ammonium ferrierite (645 g), giving a 5.4% weight loss on ignition and having the following properties: molar ratio of silica to alumina 62: 1, specific surface area 369 sq. meters per gram (P / P ₀ = 0.03), a soda content of 480 ppm and a sorption capacity for n-hexane of 7.3 g per 100 g of ammonium ferrierite, are loaded into a mixer for mixing paste Lancaster. CATAPAL ^® D alumina (91 g), giving a weight loss on ignition of 25.7%, is added to the mixer for mixing the paste. During a five minute kneading period, 152 ml of deionized water are added to the alumina / ammonium-ferrierite mixture. Then, a mixture of 6.8 g of glacial acetic acid, 7.0 g of citric acid and 152 ml of deionized water for peptization of alumina is slowly added to the alumina / ammonium-ferrierite mixture in the paste mixer. The resulting alumina / ammonium-ferrierite / acid mixture was kneaded for 10 minutes. After 15 minutes, a mixture of 0.20 g of tetraamine palladium nitrate in 153 g of deionized water was slowly added to the kneaded alumina / ammonium ferrierite / acid mixture. The resulting mixture has a zeolite to alumina ratio of 90:10 and a loss on ignition of 43.5%. The zeolite / alumina mixture is formed by extruding the mixture through a stainless steel matrix disk (1/16 ”holes) of a 2.25-inch Bonnot extruder.

The wet zeolite / alumina extrudate is dried at 125 ° C. for 16 hours. After drying, the extrudate zeolite / aluminum is again manually broken. The extrudate zeolite / alumina is calcined in a stream of air at 200 ° C for 2 hours. The temperature was raised to a maximum temperature of 500 ° C. and the zeolite / alumina extrudate was calcined for another two hours to obtain an isomerization catalyst. The isomerization catalyst is allowed to cool in a desiccator in a nitrogen atmosphere.

A stainless steel tube with an outer diameter of 1 inch, an inner diameter of 0.6 inches and a length of 26 inches is used as an isomerization reactor. The measuring channel for thermocouple entry protrudes 20 inches from the top of the stainless steel reactor tube. To load the reactor tube, it is overturned and a piece of glass wool is moved down the walls of the reactor tube through the measuring channel for introducing thermocouples and placed in the lower part of the reactor tube so that it acts as a plug for the reactor tube. Silicon carbide (20 mesh) was added to the reactor tube to a depth of approximately 6 inches. A second piece of glass wool is placed over silicon carbide. A mixture of 6.0 g of isomerization catalyst particles (6-10 mesh) and 45 g of fresh silicon carbide (60-80 mesh) are added to the reactor tube in two parts. The two-part addition distributes the isomerization catalyst evenly in the reactor tube and gives an isomerization catalyst bed of approximately 10 inches in length. A third piece of glass wool is added to the top of the catalyst in the reactor tube. Silicon carbide (20 mesh) is layered on a third piece of glass wool. A fourth piece of glass wool is placed over silicon carbide, and it serves as a plug for the bottom of the reaction tube. In order to observe the reaction temperature at various points of the reactor tube, a multi-contact thermocouple is introduced into the measuring channel for introducing the thermocouple of the reactor tube. Control the temperature above, below, and in three different parts of the catalyst bed. The reactor tube is turned over and installed in a furnace. The reactor tube is heated to an operating temperature of 280 ° C. for four hours while passing nitrogen. Upon reaching a temperature of 280 ° C., the reactor tube was held at operating temperature for another two hours to condition the isomerization catalyst.

After conditioning the isomerization catalyst with a pump, a hydrocarbon stream is fed through the reactor tube at a rate of 60 g / h. Nitrogen at a rate of 6 l / h is passed over the isomerization catalyst simultaneously with the hydrocarbon stream. The hydrocarbon stream goes into a gaseous state before reacting with an isomerization catalyst. The reactor tube operates at an outlet pressure of 20 kPa above atmospheric pressure.

Table 2 shows the mass percent of C ₈ -C ₁₀ branched olefins, C ₈ -C ₁₀ linear olefins and C ₈ -C ₁₀ paraffins in a hydrocarbon stream at 0 hours and a stream leaving the reactor tube after 24 and 48 hours of isomerization . More than 90% of the linear olefins in the hydrocarbon stream are converted to branched olefins in an isomerization reactor. During the isomerization step, a small amount of a substance boiling below C ₈ is formed by a side cleavage reaction. In addition, part of the C ₉ -C ₁₁ alcohols present in the feed undergoes dehydrogenation, giving an additional amount of olefins in the product. The average number of alkyl branches on C ₈ -C ₁₀ olefins in the product was found to be 1.0 when evaluated by ¹ H-NMR spectroscopy.

table 2 Fischer-Tropsch hydrocarbon stream composition during isomerization reaction 0 hour, wt.% 24 hours, wt.% 48 hours, wt.% C ₈ -C ₁₀ -branched olefins 0.46 33.04 33.16 C ₈ -C ₁₀ linear olefins 32.19 2,52 2.54 C ₈ -C ₁₀ paraffins 63.19 63.32 63.27 The ratio of branched to linear C ₈ -C ₁₀ olefins 0.01 13.1 13.1

Example 2: Isomerization of 1-dodecene

1-Dodecene is obtained from Shell Chemical Co. Table 3 shows the composition of 1-dodecene, determined using gas chromatography.

Table 3

Composition of 1-dodecene Wt% 1-dodecene 98.0 Other C ₁₀ -C ₁₄ olefins 1,2 <C ₁₀ hydrocarbons 0.2 > C ₁₄ hydrocarbons 0.2 Paraffins 0.4 Total C ₁₀ -C ₁₄ hydrocarbons 99.6

1-Dodecene isomerized using the same reactor tube design and the same isomerization catalyst preparation method as described in Example 1. The 1-dodecene stream is pumped through the reactor tube at a rate of 90 g / hour. Nitrogen at a rate of 6 l / h is passed over the isomerization catalyst simultaneously with a stream of 1-dodecene. The dodecene stream passes into a gaseous state before reacting with an isomerization catalyst. The reactor tube operates at an outlet pressure of 20 kPa above atmospheric pressure and at a temperature of 290 ° C.

Table 4 shows the mass percent of the molecules less than C ₁₀ , C ₁₀ -C ₁₄ , and more than C ₁₄ , in 1-dodecene after 0 hour and in the stream leaving the reactor tube after 168 and 849 hours. The linear C ₁₀ -C ₁₄ olefins are converted in 94% yield into branched C ₁₀ -C ₁₄ olefins after 168 hours of technological time. During the isomerization step, less than 3 wt.% Of a substance boiling below C ₁₀ is formed due to an adverse reaction of cleavage. The average number of alkyl branches on C ₁₀ -C ₁₄ olefins in the product was found to be 1.3 according to the results of an assessment using ¹ H-NMR spectroscopy.

Table 4 The composition of the 1-dodecene stream during the isomerization reaction 0 hour, wt.% 168 hours, wt.% 849 hours, wt.% <C ₁₀ hydrocarbons 0.2 2.5 2,4 C ₁₀ -C ₁₄ hydrocarbons 99.6 97.2 97.4 > C ₁₄ -hydrocarbons 0.2 0.3 0.2 Branched C ₁₀ -C ₁₄ olefins 0.6 93.2 93,4 Linear C ₁₀ -C ₁₄ olefins 99.0 2,8 2.9 Paraffins 1/0 2.0 1.9

Example 3: Dehydrogenation of dodecane

Dodecane is obtained from Aldrich Chemical Company and stored under nitrogen until processing. Table 5 shows the dodecane composition determined by gas chromatography.

Table 5 Dodecan Composition Wt% Dodecan 99.3 <C ₁₀ hydrocarbons <0, 1 C ₁₀ , C ₁₁ , C ₁₃ and C ₁₄ hydrocarbons <0, 6 > C ₁₄ hydrocarbons <0, 1 Other C ₁₀ -C ₁₄ olefins <0, 1

A paraffin dehydrogenation catalyst is prepared according to Example 1 (Catalyst A) of US Pat. No. 4,430,517 (Imai et al.) Entitled "Dehydrogenation Process Using a Catalytic Composition". The resulting catalyst contains 0.8 wt.% Platinum, 0.5 wt.% Tin, 2.7 wt.% Potassium and 1.3 wt.% Chlorine on a gamma-alumina substrate. The atomic ratio of potassium to palladium for this catalyst is 16.8.

The dehydrogenation catalyst is prepared by dissolving essentially pure aluminum granules in a hydrochloric acid solution. An amount of tetravalent tin chloride is added to the resulting solution to obtain a final catalyst system containing 0.5 wt.% Tin, and stirred to distribute the tin evenly throughout the mixture. Hexamethylenetetramine was added to the resulting tin mixture and the resulting tin-amine mixture was added dropwise to the oil bath so as to obtain spherical particles having an average particle diameter of about 1/6 inch. The spheres are kept, washed with an ammonia solution, dried and calcined to obtain a spherical carrier material of gamma-alumina. The resulting spheres contain approximately 0.5 wt.% Tin in the form of tin oxide. In more detail, a method for producing an alumina carrier material is described in US Pat. No. 2,620,314 (Hoesktra) under the name “Spheroidal Alumina”.

The tin (alumina) composite is brought into contact with a deionized solution of platinum hydrochloric acid and hydrochloric acid (2 wt.% Based on the weight of alumina) in a rotary dryer for 15 minutes at room temperature. The amount of platinum hydrochloric acid used is the amount needed to introduce 0.8 wt.% Platinum into the tin (alumina) composite. Then the solution is heated and purged with nitrogen to remove water to obtain a platinum-chloro-tin- (alumina) composite. The chlorine introduced is removed by heating the platinum-chloro-tin (alumina) composite to 550 ° C. and treating the composite with an air / 80 ° C. steam mixture (50:50) at a gas hourly space velocity (GHSV) of 300 h ⁻¹ . After treatment with an air / steam mixture, the platinum-tin- (alumina) composite contains less than 0.1 wt.% Chlorine.

The platinum-tin (alumina) composite is contacted with a solution of potassium nitrate in deionized water. The amount of potassium nitrate used is the amount required to introduce 2.7 wt.% Potassium into the platinum-tin (alumina) composite. Water from the platinum-tin-potassium- (alumina) composite is removed by heating the composite to 100 ° C. by blowing with dry air (GHSV 1000 h ⁻¹ ) for 0.5 h. The temperature was raised to 525 ° C and the platinum-tin-potassium- (alumina) composite was treated with a stream of hydrochloric acid (12 cc / h, 0.9 M HCl) and a mixture of air / 80 ° C steam (50:50) (GHSV 300 h ^-1 ) for introducing chlorine into the platinum-tin-potassium- (alumina) composite. The platinum-tin-potassium-chloro- (alumina) composite is dried at 525 ° C. by blowing with dry air (GHSV 1000 h ⁻¹ ). The resulting catalyst spheres have an average particle diameter of 1/16 inch, and are crushed and adjusted to a particle size of 6-20 mesh before testing.

As an isomerization reactor, a stainless steel tube with an outer diameter of 1 inch, an inner diameter of 0.6 inches and a length of 26 inches is used. The measuring channel for thermocouple entry protrudes 20 inches from the top of the stainless steel reactor tube. To load the reactor tube, it is overturned and a piece of glass wool is moved down the wall of the reactor tube through the measuring channel for thermocouple input and placed in the lower part of the reactor tube so that it acts as a plug for the reactor tube. Silicon carbide (20 mesh) was added to the reactor tube to a depth of approximately 6 inches. A second piece of glass wool is placed over silicon carbide. A mixture of 6.0 g of particles of platinum-tin catalyst on alumina (6-20 mesh) and 45 g of fresh silicon carbide (60-80 mesh) are added to the reactor tube in two parts. The two-part addition distributes the catalyst uniformly in the reactor tube and gives a catalyst bed of approximately 10 inches. A third piece of glass wool is added to the top of the catalyst in the reactor tube. Silicon carbide (20 mesh) is layered on a third piece of glass wool. A fourth piece of glass wool is placed over silicon carbide, and it serves as a plug for the bottom of the reactor tube. In order to observe the reaction temperature at various points of the reactor tube, a multi-contact thermocouple is introduced into the measuring channel for introducing the thermocouple of the reactor tube. Control the temperature above, below, and in three different parts of the catalyst bed. The reactor tube is turned over and installed in a furnace. The reactor tube is purged with nitrogen. The reactor tube is heated to an operating temperature of 425 ° C. for four hours while passing nitrogen (250 standard liters per hour). Upon reaching a temperature of 425 ° C, the reactor tube was maintained at operating temperature for another two hours. The catalyst is subjected to preliminary sulphurization by passing through the reactor tube a 1% mixture of gaseous hydrogen sulfide in hydrogen at 425 ° C for five minutes. After 5 minutes, hydrogen sulfide in a stream of hydrogen gas is switched to allow a stream of hydrogen to pass through the reactor tube.

After preliminary sulphurization of the catalyst, the reactor tube was maintained at 425 ° C for eight hours. After eight hours, the pressure in the reactor tube was raised to 25 psi using hydrogen gas. Dodecane is pumped through the reactor tube at a rate of 40 g / hr with a hydrogen rate of 125 standard liters per hour. After 4 hours, the dodecane flow is increased to 80 g / hour. After obtaining a flow rate of 80 g / h, the temperature in the reactor tube was increased to 460 ° C. After reaching an operating temperature of 460 ° C., samples were taken from the reactor tube every eight hours.

After twenty-four hours, the dodecane weight percent is 11.4 wt.%, As shown in Table 6. At a temperature of 479 ° C., the conversion of dodecane to olefins is 16 wt.% After twenty-four hours. Of the obtained olefins, 84 wt.% Are monoolefins, 4.1 wt.% Are aromatic compounds and 7.5 wt.% Are diolefins. Of the total number of olefins formed, 6% are branched olefins, according to the results of analysis using ¹ H-NMR spectroscopy.

Table 6 Test results Conversion (wt.%) After 24 hours on stream at 460 ° C 11,4 The temperature required for the conversion of 16 wt.% 479 ° C Monoolefin selectivity at 16 wt.% Conversion 84 wt.% Aromatic selectivity at 16 wt.% Conversion 4.1 wt.% Diolefin selectivity at 16 wt.% Conversion 7.5 wt.% % branched C ₁₂ olefins among all C ₁₂ olefins 6 wt.%

Example 4: Dehydrogenation-Isomerization of Dodecane

Dodecane is obtained from Aldrich Chemical Company and stored under nitrogen until processing. Table 5 shows the dodecane composition determined by gas chromatography.

The dehydrogenation-isomerization catalyst is prepared as follows. Ammonium ferrierite (645 g), giving a loss on ignition of 5.4% and having the following characteristics: molar ratio of silica to alumina 62: 1, specific surface area 369 sq. meters per gram (P / P ₀ = 0.03), a soda content of 480 ppm and a sorption capacity for n-hexane of 7.3 g per 100 g of ammonium ferrierite, are loaded into a mixer for mixing paste Lancaster. CATAPAL ^® D alumina (91 g), exhibiting a loss on ignition of 25.7%, is added to the paste mixer. During a five minute kneading period, 152 ml of deionized water are added to the alumina / ammonium-ferrierite mixture. Then, a mixture of 6.8 g of glacial acetic acid, 7.0 g of citric acid and 152 ml of deionized water for peptization of alumina is slowly added to the alumina / ammonium-ferrierite mixture in the paste mixer. The resulting alumina / ammonium-ferrierite / acid mixture was kneaded for 10 minutes. After 15 minutes, a mixture of 0.20 g of tetraamine palladium nitrate in 153 g of deionized water was slowly added to the kneaded alumina / ammonium ferrierite / acid mixture. The resulting mixture has a zeolite to alumina ratio of 90:10 and a loss on ignition of 43.5%. The zeolite / alumina mixture is formed by extruding the mixture through a stainless steel matrix disk (1/16 ”holes) of a 2.25-inch Bonnot extruder.

Six grams of the resulting zeolite / alumina mixture were impregnated with an aqueous solution of sodium (IV) hexachloroplatinate hexahydrate to introduce 0.8 wt.% Platinum into a 1/16 inch extrudate. The wet extrudate zeolite / alumina impregnated with platinum is dried at 125 ° C. for 2 hours in a stream of air. The temperature was raised to a maximum temperature of 500 ° C. and the zeolite / alumina extrudate impregnated with platinum was calcined to obtain a dehydrogenation-isomerization catalyst. The calcined catalyst is ground and adjusted to a particle size of 6-20 mesh prior to testing.

Dodecane is subjected to dehydrogenation and isomerization using the same reactor tube design as described in Example 3. The conversion of dodecane to olefins of 16.1 wt.% Is observed after twenty-four hours at 459 ° C. As shown in table 7, of the formed olefins 86 wt.% Are monoolefins, 1.2 wt.% Are aromatic compounds and 6.8 wt.% Are diolefins. Of the total amount of olefins formed, 86% are branched olefins, as determined by ¹ H-NMR spectroscopy.

Table 7 Test results Conversion (wt.%) After 24 hours on stream at 460 ° C 16.1 The temperature required for the conversion of 16 wt.% 459 ^about With Monoolefin selectivity at 16 wt.% Conversion 86 wt.%. Aromatic selectivity at 16 wt.% Conversion 1.2 wt.% Diolefin selectivity at 16 wt.% Conversion 6.8 wt.% % branched C ₁₂ olefins among all C ₁₂ olefins 86 wt.%

Example 5: Dehydrogenation-Isomerization Catalyst

The zeolite portion of the dehydrogenation-isomerization catalyst is prepared as in Example 4. Six grams of the obtained zeolite / alumina mixture are impregnated with an aqueous solution of tetraamine palladium nitrate to introduce 0.8 wt.% Palladium into a 1/16 inch extrudate.

The wet extrudate was impregnated with palladium zeolite / alumina was dried at ¹²⁵ ° C for 2 hour in a stream of air. The temperature was raised to a maximum of 500 ° C. and the palladium-impregnated zeolite / alumina extrudate was calcined to obtain a dehydrogenation-isomerization catalyst. The calcined catalyst is carefully ground and adjusted to a particle size of 6-20 mesh before testing.

Example 6: Dehydrogenation-Isomerization Catalyst

The dehydrogenation-isomerization catalyst is prepared in accordance with the process for preparing catalyst D of US Pat. No. 5,648,585 (Murray et al.) Entitled “Process for Isomerizing Linear Olefins to Isoolefins”.

Use ammonium ferrierite having a molar ratio of silica to alumina of 62: 1, a specific surface area of 369 square meters. meters per gram (P / P ₀ = 0.03), a soda content of 480 ppm and a sorption capacity for n-hexane of 7.3 g per 100 g of zeolite. The components of the catalyst are kneaded into the paste using a Lancaster paste mixer. The mixed catalytic material is extruded using a Bonnot Bonnet drum extruder. As a binder, alumina is CATAPAL ^® D from Sasol. As the extrusion additive, METHOCEL ^® F4M, hydroxypropyl methylcellulose, from Dow Chemical Company, is used.

632 g of ammonium ferrierite (3.4% loss on calcination) and CATAPAL ^® D alumina (26.2% on calcination loss) were loaded into a Lancaster paste mixer. Alumina is mixed with ferrierite for five minutes and during this time 156 ml of deionized water are added. A mixture of 6.8 g of glacial acetic acid and 156 ml of deionized water is slowly added to the paste mixing mixer to peptize the alumina. The mixture is stirred and kneaded for 10 minutes. Tetraamine platinum nitrate and tetraamine palladium tetraamine are added to the paste kneading mixer to obtain a catalyst that contains 0.25 wt.% Palladium and 0.55 wt.% Platinum. Ten grams of METHOCEL ^® F4M hydroxymethyl cellulose is added and the zeolite / alumina is kneaded for another 15 minutes. The extrudate is transferred to a Bonnot Bonnet drum extruder and extruded using a stainless steel matrix disk (1/16 inch bore). The extrudate is dried at 120 ° C for 16 hours, then calcined in air at 500 ° C for 2 hours. The calcined catalyst is carefully ground and adjusted to a size of 6-20 mesh before testing.

Claims (18)

1. The method of producing aliphatic alcohols, including:
introducing a Fischer-Tropsch hydrocarbon stream containing olefins and paraffins, said Fischer-Tropsch hydrocarbon stream containing from 5 to 80% olefins having an average carbon number of 10 to 17 and paraffins, into a hydrogenation unit, wherein the hydrogenation unit is configured to hydrogenate at least a portion of the Fischer-Tropsch hydrocarbon stream to paraffins and where at least a portion of the unreacted Fischer-Tropsch hydrocarbon stream and at least a portion of the hydrogenated olefins comfort second hydrocarbon stream;
introducing a second hydrocarbon stream into a dehydrogenation-isomerization unit, where the dehydrogenation-isomerization unit is configured to dehydrogenate at least a portion of the paraffins in the second hydrocarbon stream to olefins, and where the dehydrogenation-isomerization unit is also configured to isomerize at least , a portion of linear olefins to branched olefins in the presence of a dehydrogenation-isomerization catalyst containing a hydrogen form of a zeolite having an isotopic lattice structure of ferry and where the residence time is such that the conversion of paraffins to olefins is below 40 mol%, and where at least a portion of the unreacted components of the second hydrocarbon stream and at least a portion of the products of the dehydrogenation and isomerization reactions form a third hydrocarbon stream wherein the third hydrocarbon stream contains olefins and paraffins, and where at least a portion of the olefins in the third hydrocarbon stream are branched olefins; and
introducing at least a portion of the third hydrocarbon stream into a hydroformylation unit, wherein the hydroformylation unit is configured to hydroformylate at least a portion of the olefins in the third hydrocarbon stream to produce aliphatic alcohols with an average carbon number of 11 to 18, and where, at least a portion of the resulting aliphatic alcohols contains a branched alkyl group. 2. The method according to claim 1, where the hydrogenation unit operates in the temperature range from about 175 to 250 ° C. 3. The method according to claim 1 or 2, where the hydrogenation unit operates at a hydrogen flow rate of from about 250 to 5000 nl / l / h. 4. The method according to claim 1 or 2, where the hydrogenation unit operates in the pressure range from about 10 to 50 atmospheres. 5. The method according to claim 1 or 2, where the dehydrogenation-isomerization unit operates in the temperature range from about 300 to 500 ° C. 6. The method according to claim 1 or 2, where the installation of dehydrogenation-isomerization is made so that it works in the pressure range from about 0.010 to 15 atmospheres. 7. The method according to claim 1 or 2, where the hydroformylation unit is configured to produce approximately more than 50% aliphatic alcohols. 8. The method according to claim 1 or 2, where the hydroformylation unit operates in the range of reaction temperatures from about 100 to 300 ° C. 9. The method according to claim 1 or 2, which further includes:
forming a hydroformylation reaction stream, wherein the hydroformylation reaction stream contains at least a portion of the unreacted components of the third hydrocarbon stream and at least a portion of the obtained aliphatic alcohols;
separating aliphatic alcohols from the hydroformylation reaction stream to produce a paraffin stream and unreacted olefins and an aliphatic alcohol product stream;
the introduction of at least part of the stream of paraffins and unreacted olefins in the dehydrogenation-isomerization unit. 10. The method according to claim 9, where the introduction of at least part of the stream of paraffins and unreacted olefins in the dehydrogenation-isomerization unit includes combining at least part of the stream of paraffins and unreacted olefins with at least part of the second hydrocarbon stream obtaining a combined stream before the dehydrogenation-isomerization unit and introducing at least a portion of the combined stream into the dehydrogenation-isomerization unit. 11. The method according to claim 1 or 2, where the dehydrogenation-isomerization unit contains a dehydrogenation-isomerization catalyst, further comprising a high-melting oxide as a binder, oxidizing coke metal or a metal compound, and stimulating the dehydrogenation of paraffins metal or metal compound. 12. The method according to claim 1 or 2, where the dehydrogenation-isomerization unit includes a large number of zones, where a large number of zones include a first reaction zone and a second reaction zone, where the first reaction zone is designed to dehydrogenate at least a portion of the paraffins to olefins, and wherein the second reaction zone is designed to isomerize at least a portion of linear olefins to branched olefins, and where at least a portion of the unreacted components of the second hydrocarbon stream and at least a portion of the products dehydrogenation and isomerization reactions form a third hydrocarbon stream. 13. The method according to item 12, where the first reaction zone operates in the temperature range from about 300 to 600 ° C. 14. The method according to item 12, where the first reaction zone operates at a total reaction pressure of from about 0.01 to 25.0 atmospheres. 15. The method according to item 12, where the second reaction zone operates in the temperature range from about 200 to 500 ° C. 16. The method according to item 12, where the second reaction zone operates at a partial pressure of hydrocarbons from about 0.1 to 10 atmospheres. 17. The method according to item 12, which further includes introducing at least part of the second hydrocarbon stream exiting the first reaction zone into a heat exchanger, where the heat exchanger is configured to remove heat from part of the second hydrocarbon stream before it enters the second reaction zone. 18. The method according to claim 1 or 2, where the dehydrogenation-isomerization unit contains a catalyst in the configuration of a composite layer, where the catalyst in the form of a composite layer contains a dehydrogenation catalyst and an isomerization catalyst.

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