SYSTEMS AND METHODS FOR PRODUCING FUEL FROM A RENEWABLE FEEDSTOCK

Abstract

Methods and systems are provided for producing a fuel from a renewable feedstock. The method includes deoxygenating the renewable feedstock in a deoxygenation zone to produce hydrocarbons with normal paraffins. The hydrocarbons with normal paraffins are isomerized to produce hydrocarbons with branched paraffins. The hydrocarbons with branched paraffins are fractionated to produce a naphtha at a naphtha outlet, where the naphtha is further isomerized.

Description

TECHNICAL FIELD

The present disclosure generally relates to systems and methods for producing fuels from renewable feedstocks, and more particularly relates to systems and methods for converting renewable feedstocks into branched paraffins useful as fuel.

BACKGROUND

Many existing processes for converting renewable feedstocks into diesel fuels or jet fuels produce a naphtha stream as a co-product. The naphtha stream often includes many normal paraffin compounds, which are straight chain paraffins, that have a relatively low octane value. The octane value can be increased by isomerizing the normal paraffins into branched paraffins, because branched paraffins produce higher octane values. Increasing the octane value of the naphtha stream increases the value of the naphtha stream, and a more valuable naphtha stream increases the value of the overall process for converting renewable feedstocks into fuel.

Accordingly, it is desirable to develop methods and systems for increasing the degree of isomerization of naphtha produced as a co-product with other fuels from renewable feedstocks. In addition, it is desirable to develop methods and systems for increasing the octane value of naphtha produced from renewable feedstocks. Furthermore, other desirable features and characteristics of the present embodiment will become apparent from the subsequent detailed description and the appended claims, taken in conjunction with the accompanying drawings and this background.

BRIEF SUMMARY

A method is provided for producing fuel from renewable feedstocks. The renewable feedstock is deoxygenated in a deoxygenation zone to produce hydrocarbons with normal paraffins. The hydrocarbons with normal paraffins are isomerized to produce hydrocarbons with branched paraffins. The hydrocarbons with branched paraffins are fractionated to produce a naphtha at a naphtha outlet, where the naphtha is further isomerized.

Another method is provided for producing a fuel from a renewable feedstock. The renewable feedstock is contacted with a deoxygenation catalyst to produce hydrocarbons with normal paraffins. The hydrocarbons with normal paraffins are contacted with an isomerization catalyst to produce hydrocarbons with branched paraffins. The hydrocarbons with branched paraffins are fractionated to produce a naphtha at a naphtha outlet, and the naphtha is then isomerized.

A system is also provided for producing a fuel from a renewable feedstock. The system includes a renewable feedstock feed system coupled to a deoxygenation reaction zone. A first isomerization reaction zone is coupled to the deoxygenation reaction zone, and a fractionation zone is coupled to the first isomerization reaction zone. The fractionation zone includes a naphtha outlet, and the naphtha outlet is coupled to an isomerization reactor.

BRIEF DESCRIPTION OF THE DRAWINGS

Various embodiments will hereinafter be described in conjunction with the following figures, wherein like numerals denote like elements, and wherein:

FIG. 1 is a schematic diagram of an exemplary embodiment of a system and method for producing fuel from a renewable feedstock; and

FIG. 2 is a schematic diagram illustrating an exemplary embodiment of a system and method for fractionating and isomerizing fuel products produced from a renewable feedstock.

DETAILED DESCRIPTION

The following detailed description is merely exemplary in nature and is not intended to limit the application or uses of the embodiment described. Furthermore, there is no intention to be bound by any theory presented in the preceding technical field, background, brief summary, or the following detailed description.

Various processes for converting renewable feedstocks into fuels, especially into diesel fuel or jet fuel, also produce a naphtha co-product. The naphtha co-product is primarily hydrocarbon molecules with 5 to 8 carbon atoms and boils at a lower temperature than diesel or jet fuel. The naphtha co-product could be used for gasoline or other fuels, but it has a significant component of straight chain (normal) paraffins that have a low octane value. The octane value of the naphtha is increased by isomerizing the normal paraffins to produce branched paraffins, and the higher octane value increases the monetary value of the naphtha co-product.

Reference is now made to the exemplary embodiment illustrated in FIG. 1. A renewable feedstock 10 is processed to produce various types of fuel, such as diesel fuel, jet fuel, gasoline, liquid propane gas (LPG), etc. The term renewable feedstock 10 is meant to include feedstocks other than those derived from petroleum crude oil, and includes oils extracted from plants or animals. The renewable feedstocks 10 as contemplated herein are any of those which include glycerides or free fatty acids (FFA). Most of the glycerides will be triglycerides, but monoglycerides and diglycerides may be present and processed as well. Examples of these renewable feedstocks 10 include, but are not limited to, canola oil, corn oil, rapeseed oil, soybean oil, colza oil, tall oil, sunflower oil, hempseed oil, olive oil, linseed oil, coconut oil, castor oil, peanut oil, palm oil, mustard oil, camelina oil, pennycress oil, tallow, yellow and brown greases, lard, train oil, jatropha oil, fats in milk, fish oil, algal oil, sewage sludge, and the like. Additional examples of renewable feedstocks 10 include non-edible vegetable oils, such as oils from Madhuca indica (mahua), Pongamia pinnata, and Azadirachta indica (neem).

The glycerides and FFAs of the typical vegetable or animal fat contain aliphatic hydrocarbon chains in their structure which have about 8 to about 24 carbon atoms. The majority of the fats and oils contain high concentrations of fatty acids with 16 to 18 carbon atoms, and many types of oils contain aliphatic hydrocarbon chains within a limited range, such as 14 to 18. Only a limited number of oil types include aliphatic hydrocarbon chains covering the entire range from about 8 carbon atoms to about 24 carbon atoms, so the 8 to 24 carbon atoms range is meant to encompass mixtures of all types of oils. Co-feeds, or mixtures of renewable feedstocks 10 and petroleum derived hydrocarbons, may also be used as the feedstock. Other feedstock components that may be used, especially as a co-feed component in combination with the above listed feedstocks, include spent motor oils and industrial lubricants; used paraffin waxes; liquid derived from the gasification of coal, biomass, or natural gas followed by a downstream liquefaction step such as Fischer-Tropsch technology; liquids derived from depolymerization (thermal or chemical) of waste plastics such as polypropylene, high density polyethylene, and low density polyethylene; and other synthetic oils generated as byproducts from petrochemical and chemical processes. Mixtures of the above feedstocks may also be used as co-feed components. One advantage of using a co-feed component is the transformation of what may have been a waste product into a valuable co-feed component to the current process.

The renewable feedstock 10 is stored and delivered for processing by a renewable feedstock feed system 12. In an exemplary embodiment, the renewable feedstock feed system 12 includes a renewable feedstock storage tank 14, renewable feedstock pump 16, and associated piping. The renewable feedstock feed system 12 delivers a renewable feedstock feed stream 18 for further processing. Other embodiments of the renewable feedstock feed system 12 exist, such as a pipeline from a different source, and a pressurized renewable feedstock storage tank 14 without a renewable feedstock pump 16.

Many renewable feedstocks 10 that can be used herein contain a variety of impurities. For example, tall oil is a byproduct of the wood processing industry, and includes esters and rosin acids in addition to FFAs. Rosin acids are cyclic carboxylic acids. The renewable feedstocks 10 may also contain contaminants such as alkali metals, (e.g. sodium and potassium), phosphorous, various solids, water, and detergents. In some embodiments, the renewable feedstock 10 is pre-cleaned in an optional pre-cleaning zone 20 to improve downstream processing operations, and several different types of pre-cleaning are possible. For example, the pre-cleaning zone 20 may be configured to provide a mild acid wash by contact with dilute sulfuric, nitric, citric, phosphoric, or hydrochloric acid in a reactor. The acid wash can be a continuous process or a batch process, and the dilute acid contact can be at ambient temperature and atmospheric pressure. Other possible pre-cleaning steps include, but are not limited to, contacting the renewable feedstock 10 with an ion exchange resin such as Amberlyst®-15, subjecting the renewable feedstock 10 to a caustic treatment, bleaching the renewable feedstock 10 with an adsorbent, filtration, solvent extraction, hydro processing, or combinations of the above.

In some embodiments, a sulfiding agent 22 is added to the renewable feedstock 10. Several reactors described more fully below use catalysts of various types, and one or more of these catalysts can be used in a sulfided state in various embodiments. Sulfur is added to the process to maintain the catalysts in the sulfided state. The sulfiding agent 22 is added at a sulfiding agent inlet 24. The sulfur is measured as elemental sulfur, regardless of the compound containing the sulfur, and can be added in many forms. For example, suitable sulfiding agents 22 include, but are not limited to, dimethyl disulfide, dibutyl disulfide, and hydrogen sulfide. The sulfur may be obtained from various sources, such as part of a hydrogen stream from a hydrocracking unit or hydro treating unit, or sulfur compounds removed from kerosene or diesel, and disulfide oils removed from sweetening units such as Merox® units. A deoxygenation catalyst is described more fully below, and sulfur concentrations of less than 2,000 ppm are typically sufficient to maintain the deoxygenation catalyst and the other catalysts described below in a sulfided state. FIG. 1 illustrates adding the sulfiding agent 22 to the renewable feedstock feed stream 18, but other embodiments are possible. For example, some renewable feedstocks 10 contain sufficient sulfur to maintain the catalysts in a sulfided state. Sulfur can also be added to the renewable feedstock storage tank 14, the reactors containing the catalysts, or other locations.

In an exemplary embodiment, a recycle hydrogen stream 80 (described more fully below) is added to the renewable feedstock feed stream 18 and flows downstream to a guard bed 26. A portion of a hot separator bottoms stream 50 (described more fully below) is also added to the renewable feedstock feed stream 18 before entry into the guard bed 26. The guard bed 26 removes metals from the renewable feedstock 10 by contacting the renewable feedstock feed stream 18 with a guard bed catalyst 28 at pretreatment conditions. The guard bed catalyst 28 may initiate a deoxygenation reaction of the renewable feedstock feed stream 18 to some degree, as described more fully below. In some embodiments, the guard bed catalyst 28 is alumina, either with or without demetallation catalysts such as nickel or cobalt, but other guard bed catalysts 28 are also possible. The guard bed 26 is operated at a temperature from about 40° C. to about 400° C., for example from about 150° C. to about 300° C. Operating pressures for the guard bed 26 are from about 690 kilopascals (kPa) absolute (100 pounds per square inch absolute (psia)) to about 13,800 kPa absolute (2,000 psia), for example from about 1,380 kPa absolute (200 psia) to about 6,900 kPa absolute (1,000 psia). A portion of the hot separator bottoms stream 50 may be added at various locations in the guard bed 26 to aid in temperature control, hydrogen solubility, or other purposes, but in other embodiments different streams or no streams are added at side locations in the guard bed 26.

After the optional guard bed 26, a guard bed effluent 30 flows downstream to a deoxygenation reaction zone 40 including one or more catalyst beds in one or more reactors. In the deoxygenation reaction zone 40, the guard bed effluent 30 is contacted with a deoxygenation catalyst 42 (sometimes referred to as a hydrotreating catalyst) in the presence of hydrogen at deoxygenation conditions. The hydrogen for this reaction is provided from the recycle hydrogen stream 80 added to the renewable feedstock feed stream 18. Under these conditions, the olefinic or unsaturated portions of n-paraffinic chains are hydrogenated. Additionally, any deoxygenation reactions that did not take place in the guard bed 26 are completed in the deoxygenation reaction zone 40. In some embodiments, a portion of the hot separator bottoms stream 50 is added at various locations in the deoxygenation reaction zone 40 to aid in temperature control, hydrogen solubility, and other purposes. In other embodiments, streams other than the hot separator bottoms stream 50 (or even no streams) are added at side locations in the deoxygenation reaction zone 40. A deoxygenation effluent 44 exits the deoxygenation reaction zone 40.

Deoxygenation catalysts 42 are any of those well known in the art, such as nickel, nickel/molybdenum, or cobalt/molybdenum dispersed on a high surface area support. Other deoxygenation catalysts 42 include one or more noble metal catalytic elements dispersed on a high surface area support. Non-limiting examples of noble metals include platinum (Pt) and/or palladium (Pd). Deoxygenation conditions include a temperature of about 40 degrees centigrade (° C.) to about 400° C., and a pressure of about 690 kilopascals (kPa) absolute (100 psia) to about 13,800 kPa absolute (2,000 psia). In another embodiment the deoxygenation conditions include a temperature of about 200° C. to about 300° C., and a pressure of about 1,380 kPa absolute (200 psia) to about 6,900 kPa absolute (1,000 psia). Other operating conditions for the deoxygenation reaction zone 40 can also be used. A sulfiding agent 22, such as from the sulfiding agent inlet 24 or from the renewable feedstock 10, maintains the deoxygenation catalyst 42 in a sulfided state.

The deoxygenation catalysts 42 discussed above are also capable of catalyzing decarboxylation, decarbonylation and/or hydrodeoxygenation of the renewable feedstock 10 to remove oxygen. Decarboxylation, decarbonylation, and hydrodeoxygenation are herein collectively referred to as “deoxygenation reactions”, and the deoxygenation reactions and the olefin hydrogenation reactions simultaneously occur in the deoxygenation reaction zone 40. Deoxygenation conditions include a relatively low pressure of about 3,450 kPa (500 psia) to about 6,900 kPa (1,000 psia), a temperature of about 200° C. to about 400° C., and a liquid hourly space velocity of about 0.2 to about 10 hr⁻¹. In another embodiment the deoxygenation conditions include the same relatively low pressure of about 3,450 kPa (500 psia) to about 6,900 kPa (1,000 psia), a temperature of about 290° C. to about 350° C., and a liquid hourly space velocity of about 1 to about 4 hr⁻¹.

Deoxygenation is an exothermic reaction, so the temperature in the deoxygenation reaction zone 40 increases as the hydrocarbons from the renewable feedstock 10 pass through. Decarboxylation and hydrodeoxygenation reactions begin to occur as the temperature increases. The rate of the deoxygenation reactions increases from the front of the bed to the back of the bed as the temperature increases. The deoxygenation reaction zone 40 can include one or more reactors in series, and can also include parallel reactors or sets of reactors.

The hydrodeoxygenation reaction consumes hydrogen and produces water as a byproduct, while the decarbonylation and decarboxylation reactions produce carbon monoxide (CO) or carbon dioxide (CO₂) without consuming hydrogen. However, hydrogen is present for all the reactions in the deoxygenation reaction zone 40, regardless of whether the reaction consumes hydrogen or not. The product from the deoxygenation reactions includes a liquid portion and a gaseous portion. The liquid portion present in the deoxygenation effluent 44 includes hydrocarbon compounds that are largely normal paraffin compounds (n-paraffins) having a high cetane number. The gaseous portion includes hydrogen, carbon dioxide (CO₂), carbon monoxide (CO), water vapor, propane, and perhaps sulfur components such as hydrogen sulfide. It is possible to separate and collect the liquid portion (the hydrocarbons including n-paraffins) as a diesel fuel product without further reactions. However, in most climates, at least a portion of the liquid n-paraffins can be isomerized to produce branched paraffins, which improves the cold flow properties of the fuel.

In an exemplary embodiment, the deoxygenation effluent 44 passes to an optional hot separator 46 downstream from the deoxygenation reaction zone 40. One purpose of the hot separator 46 is to separate at least some of the gaseous portion from the liquid portion of the deoxygenation effluent 44. Much of the gaseous portion, including the recovered hydrogen, exits the hot separator 46 in a hot separator overhead stream 48, and the liquid portion exits the hot separator in a hot separator bottoms stream 50. The separated hydrogen is recycled back to the deoxygenation reaction zone 40 in some embodiments, as described more fully below. The liquid hydrocarbons including the n-paraffins exit the hot separator 46 in the hot separator bottoms stream 50.

In some embodiments, water, CO, CO₂, and any ammonia or hydrogen sulfide are stripped in the hot separator 46 using hydrogen. In some embodiments (not shown), additional hydrogen is used as the stripping gas, but other gases could also be used. The temperature is controlled to achieve the desired separation, and the pressure can be maintained at approximately the same pressure as the deoxygenation reaction zone 40 and the isomerization reaction zone (described below) to minimize both investment and operation costs. Energy is required to change the temperature or pressure, which increases operating costs, and additional equipment is needed to enable the process to change the temperature of pressure, which increases the investment cost. The hot separator 46 may be operated at conditions ranging from a pressure of about 690 kPa absolute (100 psia) to about 13,800 kPa absolute (2,000 psia), and a temperature of about 40° C. to about 350° C. In another embodiment, the hot separator 46 may be operated at conditions ranging from a pressure of about 1,380 kPa absolute (200 psia) to about 6,900 kPa absolute (1,000 psia), or about 2,410 kPa absolute (350 psia) to about 4,880 kPa absolute (650 psia), and a temperature of about 50° C. to about 350° C.

The paraffinic components of the hot separator bottoms stream 50 are primarily n-paraffins which range from about 8 to about 24 carbon atoms depending on the type of renewable feedstock 10 used. Different renewable feedstocks 10 will result in different distributions of paraffins. The hot separator bottoms stream 50 is divided and transferred to various locations in different embodiments. A portion of the hot separator bottoms stream 50 may be recycled and added to the guard bed 26 at various locations, and to the deoxygenation reaction zone 40 at various locations, as described above. In alternate embodiments, other streams or no streams are recycled in place of the hot separator bottoms stream 50.

In an exemplary embodiment, the hot separator bottoms stream 50 also flows to an enhanced hot separator 52 to further separate the gaseous and liquid components of the deoxygenation effluent 44. Additional gases are removed from the liquid hydrocarbons with the n-paraffins, and the gases are vented in an enhanced hot separator overhead stream 54, which is combined with the hot separator overhead stream 48. The enhanced hot separator 52 operates at similar conditions as the hot separator 46. The enhanced hot separator operating conditions range from a pressure of about 690 kPa absolute (100 psia) to about 13,800 kPa absolute (2,000 psia), and a temperature of about 40° C. to about 350° C. In another embodiment, the enhanced hot separator 52 may be operated at conditions ranging from a pressure of about 1,380 kPa absolute (200 psia) to about 6,900 kPa absolute (1,000 psia), or about 2,410 kPa absolute (350 psia) to about 4,880 kPa absolute (650 psia), and a temperature of about 50° C. to about 350° C.

An enhanced hot separator bottoms stream 56 flows from the enhanced hot separator 52 downstream to a first isomerization reaction zone 60. The enhanced hot separator bottoms stream 56 is primarily made up of the liquid hydrocarbons, including the n-paraffins, from the deoxygenation reaction zone 40. Fresh hydrogen is added to the enhanced hot separator bottoms stream 56 from a hydrogen feed line 36, so additional hydrogen is fed to the first isomerization reaction zone 60. In other embodiments, the hydrogen could be fed to the first isomerization reaction zone 60 in other manners, such as from a feed line piped directly into a reactor in the first isomerization reaction zone 60.

Isomerization can be carried out in a separate bed of the same reactor used in the deoxygenation reaction zone 40, or the isomerization can be carried out in a separate isomerization reactor 58. For ease of description, the following will address the embodiments where a separate reaction zone is employed for the first isomerization reaction zone 60. In an exemplary embodiment, the first isomerization reaction zone 60 includes an isomerization catalyst 62 positioned within an isomerization reactor 58, and is operated at isomerization conditions. The hydrocarbons with the n-paraffins in the enhanced hot separator bottoms stream 56 are contacted with the isomerization catalyst 62 in the presence of hydrogen to convert at least some of the n-paraffins into branched paraffins. Only minimal branching is required to overcome the poor cold-flow characteristics of the n-paraffins used in diesel or jet fuel. In some embodiments, the predominant isomerized paraffin product is a mono-branched hydrocarbon, because process conditions that produce significant branching also increase the risk of excessive cracking that reduces the yield of diesel or jet fuel. The hydrocarbons used in diesel and jet fuel generally have more carbons than the hydrocarbons used in gasoline, on average, and have a higher average boiling point. Besides improving the cold flow properties of diesel fuel, branched paraffins also increase the octane rating of gasoline fuels.

An isomerization effluent 64, which exits the first isomerization reaction zone 60, is a hydrocarbon stream rich in branched paraffins. By the term “rich” it is meant that the isomerization effluent 64 has a greater concentration of branched paraffins than the stream entering the first isomerization reaction zone 60, and in some embodiments includes greater than 50 mass percent branched paraffins. The isomerization effluent 64 may contain 70, 80, or 90 mass percent branched paraffins in some embodiments, but lower concentrations of branched paraffins are present in other embodiments. The degree of isomerization can be changed by adjusting the isomerization conditions. For example, a lower reactor temperature will decrease the degree of isomerization, and also decrease the degree of cracking in the first isomerization reaction zone 60.

The isomerization of the n-paraffins can be accomplished by using a variety of suitable catalysts. The first isomerization reaction zone 60 includes one or more beds of isomerization catalyst 62, and the catalyst beds can be in series and/or parallel. A single isomerization reactor 58 may include one or more catalyst beds, so the first isomerization reaction zone 60 can also include one or more isomerization reactors 58. In some embodiments, the first isomerization reaction zone 60 is operated in a co-current mode of operation. Fixed bed trickle down flow or fixed bed liquid upward flow modes are both suitable. In some embodiments, the isomerization catalyst 62 is not sulfided, so no sulfiding agents are added to streams entering the first isomerization reaction zone 60 downstream from the deoxygenation reaction zone 40. In alternate embodiments, the isomerization catalyst is sulfided.

Suitable isomerization catalysts 62 include a metal of Group VIII (IUPAC 8-10) of the Periodic Table and a support material. Suitable Group VIII metals include platinum and palladium, each of which may be used alone or in combination. The support material may be amorphous or crystalline, and many different support materials can be used. Suitable support materials include, but are not limited to, amorphous alumina, amorphous silica-alumina, ferrierite, metal aluminumsilicophosphates, laumontite, cancrinite, offretite, the hydrogen form of stillbite, the magnesium or calcium form of mordenite, and the magnesium or calcium form of partheite, each of which may be used alone or in combination. Many natural zeolites, such as ferrierite, that have an initially reduced pore size can be converted to forms suitable for olefin skeletal isomerization by removing associated alkali metals or alkaline earth metals by ammonium ion exchange and calcination to produce a substantial hydrogen form. The isomerization catalyst 62 may also include one or more modifiers, such as those selected from the group of lanthanum, cerium, praseodymium, neodymium, samarium, gadolinium, terbium, and mixtures thereof

The isomerization reaction occurs when hydrocarbons pass through the isomerization catalyst 62 at isomerization conditions. Isomerization conditions include a temperature of about 150° C. to about 420° C. and a pressure of about 1,720 kPa absolute (250 psia) to about 4,720 kPa absolute (700 psia). In another embodiment, the isomerization conditions include a temperature of about 300° C. to about 360° C. and a pressure of about 2,400 kPa absolute (350 psia) to about 3,800 kPa absolute (550 psia). Other operating conditions for the first isomerization reaction zone 60 can also be used.

The hydrocarbons with the branched paraffins in the isomerization effluent 64 are processed through one or more separation steps to obtain a hydrocarbon stream useful as a fuel, and the separation steps vary in different embodiments. The isomerization effluent 64 includes both a liquid component and a gaseous component, various portions of which can be recycled, so multiple separation steps may be employed. For example, in some embodiments the isomerization effluent 64 is separated in an isomerization effluent separator 66 positioned downstream from the first isomerization reaction zone 60. Hydrogen exits the isomerization effluent separator 66 in an isomerization effluent separator overhead stream 68, and the liquid portion exits in an isomerization effluent separator bottoms stream 70. The isomerization effluent separator overhead stream 68 is fed to the enhanced hot separator 52 in some embodiments, so the gaseous portions are combined with the gases in the enhanced hot separator overhead stream 54. In other embodiments (not shown), the isomerization effluent separator overhead stream 68 bypasses the enhanced hot separator 52 and is eventually used as recycled hydrogen or processed in other ways.

Suitable operating conditions of the isomerization effluent separator 66 include, for example, a temperature of about 280° C. to about 360° C. and a pressure of about 4,100 kPa absolute (600 psia), but other operating conditions are also possible. If there is a low concentration of carbon oxides, or the carbon oxides are removed, the hydrogen may be directly recycled and re-used in the process. Hydrogen is a reactant in the deoxygenation reaction zone 40 and the first isomerization reaction zone 60, and different renewable feedstocks 10 will consume different amounts of hydrogen. Additional hydrogen can be added for feeds that consume more hydrogen. Furthermore, at least a portion of the isomerization effluent separator bottoms stream 70 can be recycled to the first isomerization reaction zone 60 (not shown) to increase the degree of isomerization, to aid in temperature control, or for other purposes.

The remainder of the isomerization effluent separator bottoms stream 70 still has liquid and gaseous components and can be cooled by various techniques, such as air cooling or water cooling. The liquid portion of the isomerization effluent separator bottoms stream 70 is hydrocarbons, including the branched paraffins, as well as some n-paraffins that were not isomerized into branched paraffins. After cooling, the isomerization effluent separator bottoms stream 70 is passed to a cold separator 72 where the liquid component is separated from the gaseous component. The hot separator overhead stream 48 and the enhanced hot separator overhead stream 54 are also fed to the cold separator 72, and can be combined with the isomerization effluent separator bottoms stream 70 upstream from the cold separator 72. Suitable operating conditions of the cold separator 72 include, for example, a temperature of about 40° C. to about 60° C. (about 100° F. to about 140° F.) and a pressure of about 3,800 kPa absolute to about 5,300 kPa absolute (about 550 to about 770 psia), but other operating conditions are also possible. A water byproduct stream is also separated in the cold separator 72 (not shown). A cold separator overhead stream 74 and a cold separator bottoms stream 76 exit the cold separator 72.

The cold separator overhead stream 74, or the gaseous components separated in the cold separator 72, is mostly hydrogen and the carbon dioxide from the decarboxylation reaction. Other components such as CO, propane, and hydrogen sulfide or other sulfur containing components may be present as well. Water, CO, and CO₂ can negatively impact the catalyst performance in the first isomerization reaction zone 60. It is desirable to recycle the hydrogen, but if the CO₂ and other components are not removed, their concentrations can build up and negatively affect the operation of the first isomerization reaction zone 60. A recovery gas cleaner 78 can be used to increase the purity of the cold separator overhead stream 74. The carbon dioxide can be removed from the hydrogen by several different processes, including but not limited to absorption with an amine, reaction with a hot carbonate solution, pressure swing absorption, etc. If desired, essentially pure carbon dioxide can be recovered by regenerating the spent absorption media. A sulfur containing component, such as hydrogen sulfide, may also be present. The sulfur containing component may be used to help control the relative amounts of the decarboxylation reaction and the hydrogenation reaction in the deoxygenation reaction zone 40. The amount of sulfur is generally controlled, so the sulfur is also removed before the hydrogen is recycled. Various methods can be used, such as absorption with an amine or a caustic wash, and the carbon dioxide and sulfur containing components (as well as other components) are removed in a single separation step in some embodiments.

A recycle hydrogen stream 80 exits the recovery gas cleaner 78 after the impurities have been removed. A recycle hydrogen compressor 82 urges the hydrogen back into the process. As discussed above, the recycle hydrogen stream 80 may be fed into the renewable feedstock feed stream 18, but the recycle hydrogen stream 80 could be routed into the process in other locations as well, such as routed directly into the reactors of the deoxygenation reaction zone 40 or the first isomerization reaction zone 60. The recycle hydrogen stream 80 supplies the hydrogen for the guard bed 26 and the deoxygenation reaction zone 40, as discussed above.

The cold separator bottoms stream 76, or the liquid component separated in the cold separator 72, contains the liquid hydrocarbons with the branched paraffins useful as jet fuel and/or diesel fuel, as well as smaller amounts of naphtha, liquid propane gas (LPG), and other hydrocarbons. The cold separator bottoms stream 76 may be recovered as diesel boiling range fuel or it may be further purified in a fractionation zone 84 that fractionates the various components of the cold separator bottoms stream. In one embodiment, the fractionation zone 84 includes a product stripper 86 or a product fractionator (not shown) that can be operated, for example, with a vapor temperature of from about 20° C. to about 200° C. and a pressure from about 0 kPa (0 psia) to about 1,380 kPa absolute (200 psia) at the overhead of the product stripper 86. In alternate embodiments, the fractionation zone 84 includes a plurality of fractionators and/or separators to divide the cold separator bottoms stream 76 into various fractions. The fractionation zone 84 separates the cold separator bottoms stream 76 into a fractionation zone overhead stream 88, a naphtha product that exits the fractionation zone 84 at a naphtha outlet 92, and a fractionation zone bottoms stream 94. The naphtha outlet 92 is split into a naphtha fraction 90 that is collected as a product, and a naphtha reisomerization stream 93.

The fractionation zone overhead stream 88 includes LPG and lighter hydrocarbons, such as ethane or methane, and it may include butanes. The fractionation zone overhead stream 88 can be further fractionated and sold as a product, used as a fuel gas, or used in other processes such as the feed to a hydrogen production facility, a co-feed to a reforming process, or a fuel blending component. The fractionation zone bottoms stream 94 can be used a diesel range fuel or further fractionated and used as a jet fuel. The naphtha fraction 90 includes hydrocarbons with about 5 to 8 carbon atoms, and boils from about 20° C. to about 150° C., where the hydrocarbons are primarily a mixture of n-paraffins and branched paraffins. In some embodiments, the naphtha is lightly isomerized after making one pass through an isomerization reactor 58, so it includes relatively few branched paraffins. The naphtha fraction 90 can be used as a component in gasoline, but it has an octane value of about 60 to about 70 after a single pass through the isomerization reactor 58, so a higher octane value would increase the value of the naphtha fraction 90 for use in gasoline. Most gasoline sold commercially has an octane value of about 85 to about 95. The octane value can be increased by converting n-paraffins into branched paraffins.

In an exemplary embodiment, some of the naphtha product from the naphtha outlet 92 is further isomerized to convert n-paraffins into branched paraffins by routing the naphtha reisomerization stream 93 back to the first isomerization reaction zone 60. Some of the naphtha product is removed from the process in a naphtha fraction 90 to prevent the naphtha from building up in the system. The isomerization catalyst 62 in the first isomerization reaction zone 60 will crack some of the hydrocarbons in the naphtha into smaller molecules, which decreases the yield of the final naphtha fraction 90. However, cracking of the hydrocarbons in the naphtha is minimized by reducing the contact time with the isomerization catalyst 62 in the first isomerization reaction zone 60. The naphtha reisomerization stream 93 may be added to the first isomerization reaction zone 60 by coupling the naphtha outlet 92 to a side inlet 96 of an isomerization reactor 58 in the first isomerization reaction zone 60, where the side inlet 96 is positioned with some of the catalyst bed upstream from the side inlet 96 and some of the catalyst bed downstream from the side inlet 96. The position of the side inlet 96 can be adjusted to optimize the degree of isomerization of the naphtha with the degree of cracking, and in some embodiments the naphtha reisomerization stream 93 is coupled to the inlet of the isomerization reactor 58 and contacted with the entire isomerization catalyst bed. A side inlet 96 configured so the naphtha bypasses some of the isomerization catalyst 62 also minimizes any dilution effect by the naphtha on the isomerization of the hydrocarbons with n-paraffins in the enhanced hot separator bottoms stream 56.

Reference is now made to the exemplary embodiment illustrated in FIG. 2, which begins with the cold separator bottoms stream 76. In this embodiment, the fractionation zone 84 includes a product stripper 86 with a fractionation zone overhead stream 88 and a fractionation zone bottoms stream 94. The fractionation zone overhead stream 88 is fed into a light gas separator 98. A lean gas stream 100 exits the light gas separator as a gas, and a light gas separator bottoms stream 102 exits as a liquid. The light gas separator bottoms stream 102 from the light gas separator 98 includes the LPG 104 and the hydrocarbons in the naphtha fraction 118. The LPG 104 and hydrocarbons in the naphtha fraction 118 (prior to isomerization) are further separated in a debutanizer 106 that produces the LPG 104 as an overhead stream and the naphtha reisomerization stream 93 as a bottom stream. The naphtha reisomerization stream 93 exits the debutanizer 106 at the naphtha outlet 92. The debutanizer 106 can be operated, for example, at a vapor temperature of about 20° C. to about 200° C. and a pressure from about 0 to about 2,760 kPa absolute (0 to 400 psia) at the debutanizer overhead, but other conditions are also possible.

The naphtha outlet 92 from the debutanizer 106 is coupled to an isomerization reactor 114 in a second isomerization reaction zone 110 to further isomerize the paraffins in the naphtha fraction 118. In some embodiments, the second isomerization reaction zone 110 includes an isomerization catalyst 116 and operates at isomerization conditions. The second isomerization reaction zone 110 can be operated to match the feed from the naphtha outlet 92, and a suitable isomerization catalyst 116 and isomerization conditions can be used. In an exemplary embodiment, the isomerization catalyst 116 includes about 0.01 to about 3 weight percent of a metal on an inorganic oxide carrier, and includes a halide as a promoter. Suitable inorganic oxide carriers include alumina, silica, zirconia, magnesia, thoria, and combinations thereof, but other carriers can also be used. Suitable metals include Ruthenium, Rhodium, Palladium, Osmium, Iridium, and Platinum, and the weight percent is determined based on the weight of the metal, regardless of the form of the metal on the carrier. The halide promoter is present at about 0.1 to about 10 weight percent, and includes chlorides or other halides. Suitable isomerization conditions include a temperature from about 120° C. to about 200° C. (about 250° F. to about 400° F.), and pressures from about 2,400 kPa to about 3,800 kPa (about 350 PSIG to about 550 PSIG).

The second isomerization reaction zone 110 can be used in place of, or in conjunction with, a naphtha recycle through the first isomerization reaction zone. The naphtha reisomerization stream 93 is the primary feed to the second isomerization reaction zone 110, so the size of the isomerization reactor 114 and catalyst bed, the quantity of isomerization catalyst 116 used, and the isomerization conditions can be optimized for the naphtha reisomerization stream 93. A second isomerization reaction zone hydrogen line 112 can be used to introduce hydrogen for the isomerization reaction. The naphtha fraction 118 then exits the second isomerization reaction zone 110 with a higher level of branched paraffins than the feed to the second isomerization reaction zone 110. An optional separator (not shown) can be installed downstream from the second isomerization reaction zone 110 to vent hydrogen and light gases produced by cracking in the isomerization reactor 114, and the vented hydrogen can be reused in a similar manner to the hydrogen collected in the hot separator overhead stream.

Reference is now made to FIG. 1 again. The exemplary embodiments described above include many optional processes, or processes that can be modified or arranged in different manners. In a very simplified form, the renewable feedstock feed system 12 is coupled to the deoxygenation reaction zone 40, because the renewable feedstock 10 flows to the deoxygenation reaction zone 40. The deoxygenation reaction zone 40 is likewise coupled to the first isomerization reaction zone 60, which is coupled to the fractionation zone 84, even though several vessels or processes are positioned between the different zones. The naphtha is recovered from the fractionation zone 84 and re-isomerized to increase the concentration of branched paraffins. Several vessels and steps are used to recover and reuse hydrogen throughout the manufacturing process.

It should be appreciated that the embodiment or embodiments illustrated are only examples, and are not intended to limit the scope, applicability, or configuration of the application in any way. Rather, the foregoing detailed description will provide those skilled in the art with a convenient road map for implementing one or more embodiments, it being understood that various changes may be made in the function and arrangement of elements described without departing from the scope as set forth in the appended claims.

Claims

1. A method of producing fuel from a renewable feedstock, the method comprising the steps of:

deoxygenating the renewable feedstock in a deoxygenation reaction zone to produce hydrocarbons comprising normal paraffins;

isomerizing the hydrocarbons comprising normal paraffins to produce hydrocarbons comprising branched paraffins;

fractionating the hydrocarbons comprising branched paraffins to produce a naphtha at a naphtha outlet; and

isomerizing the naphtha from the naphtha outlet.

2. The method of claim 1 wherein isomerizing the hydrocarbons comprising normal paraffins further comprise isomerizing the hydrocarbons comprising normal paraffins in a first isomerization reaction zone at isomerization conditions; and wherein isomerizing the naphtha further comprises isomerizing the naphtha in the first isomerization reaction zone.

3. The method of claim 2 wherein isomerizing the naphtha further comprises adding the naphtha to the first isomerization reaction zone such that the naphtha bypasses a portion of an isomerization catalyst positioned within the first isomerization reaction zone.

4. The method of claim 2 wherein isomerizing the naphtha further comprises isomerizing the naphtha in a second isomerization reaction zone.

5. The method of claim 1 wherein isomerizing the naphtha further comprises isomerizing the naphtha in a second isomerization reaction zone.

6. The method of claim 1 wherein deoxygenating the renewable feedstock further comprises deoxygenating the renewable feedstock wherein the renewable feedstock comprises glycerides or free fatty acids.

7. The method of claim 1 wherein deoxygenating the renewable feedstock further comprises deoxygenating the renewable feedstock wherein the renewable feedstock comprises oil extracted from a plant or an animal.

8. The method of claim 1 further comprising sulfiding a deoxygenation catalyst in the deoxygenation reaction zone.

9. The method of claim 1 further comprising:

contacting the renewable feedstock with a guard bed catalyst at pretreatment conditions.

10. The method of claim 1 further comprising:

pre-cleaning the renewable feedstock in a pre-cleaning zone.

11. A method of producing fuel from a renewable feedstock, the method comprising the steps of:

contacting the renewable feedstock with a deoxygenation catalyst to produce hydrocarbons comprising normal paraffins;

contacting the hydrocarbons comprising normal paraffins with an isomerization catalyst to produce hydrocarbons comprising branched paraffins;

fractionating the hydrocarbons comprising branched paraffins to produce a naphtha at a naphtha outlet; and

isomerizing the naphtha from the naphtha outlet.

12. The method of claim 11 wherein contacting the hydrocarbons comprising normal paraffins with the isomerization catalyst further comprises contacting the hydrocarbons comprising normal paraffins with the isomerization catalyst wherein the isomerization catalyst is within a first isomerization reaction zone; and wherein isomerizing the naphtha further comprises contacting the naphtha with the isomerization catalyst in the first isomerization reaction zone.

13. The method of claim 12 wherein isomerizing the naphtha further comprises adding the naphtha to an isomerization reactor at a side inlet of the isomerization reactor, wherein the isomerization catalyst is positioned within the isomerization reactor and wherein the side inlet is positioned such that the naphtha bypasses some of the isomerization catalyst within the isomerization reactor.

14. The method of claim 12 wherein isomerizing the naphtha further comprises contacting the naphtha with the isomerization catalyst in a second isomerization reaction zone.

15. The method of claim 11 wherein contacting the hydrocarbons comprising normal paraffins with the isomerization catalyst further comprises contacting the hydrocarbons comprising normal paraffins with the isomerization catalyst wherein the isomerization catalyst is within a first isomerization reaction zone; and wherein isomerizing the naphtha further comprises isomerizing the naphtha in a second isomerization reaction zone different than the first isomerization reaction zone.

16. The method of claim 11 wherein contacting the renewable feedstock with the deoxygenation catalyst further comprises contacting the renewable feedstock with the deoxygenation catalyst wherein the renewable feedstock comprises glycerides or free fatty acids.

17. The method of claim 11 wherein contacting the renewable feedstock with the deoxygenation catalyst further comprises contacting the renewable feedstock with the deoxygenation catalyst wherein the renewable feedstock comprises oil extracted from a plant or an animal.

18. The method of claim 11 further comprising sulfiding the deoxygenation catalyst.

19. The method of claim 1 further comprising:

pre-cleaning the renewable feedstock in a pre-cleaning zone.

20. A system for producing fuel from a renewable feedstock comprising;

a renewable feedstock feed system;

a deoxygenation reaction zone coupled to the renewable feedstock feed system;

a first isomerization reaction zone coupled to the deoxygenation reaction zone;

a fractionation zone coupled to the first isomerization reaction zone, wherein the fractionation zone comprises a naphtha outlet; and

an isomerization reactor, wherein the naphtha outlet is coupled to the isomerization reactor.