Patent Application
20180216010
This invention relates to a process for producing diesel boiling range fuel and jet fuel from renewable feedstocks such as the triglycerides and free fatty acids found in materials such as plant oils, fish oils, animal fats, and greases. Hydrogenation and deoxygenation are performed in one or more reactors without use of expensive purified hydrogen. One or more sulfur compounds may be present in the reaction mixture. A vapor stream is separated from the reaction zone effluent, and carbon dioxide is separated from the vapor stream. Optionally, a sulfur-containing stream may also be separated.
As the demand for diesel boiling range fuel and jet fuel increases worldwide there is increasing interest in sources other than crude oil for producing diesel boiling range fuel and jet fuel. One such renewable source is what has been termed biorenewable sources. These renewable sources include, but are not limited to, plant oils such as corn, rapeseed, canola, soybean and algal oils, animal fats such as inedible tallow, fish oils and various waste streams such as yellow and brown greases and sewage sludge. The common feature of these sources is that they are composed of glycerides and Free Fatty Acids (FFA). Both of these classes of compounds contain aliphatic carbon chains having from about 8 to about 24 carbon atoms. The aliphatic carbon chains in the glycerides or FFAs can be fully saturated, or mono, di or poly-unsaturated.
There are reports in the art disclosing the production of hydrocarbons from oils. For example, U.S. Pat. No. 4,300,009 discloses the use of crystalline aluminosilicate zeolites to convert plant oils such as corn oil to hydrocarbons such as gasoline and chemicals such as para-xylene. U.S. Pat. No. 4,992,605 discloses the production of hydrocarbon products in the diesel boiling range by hydroprocessing vegetable oils such as canola or sunflower oil. Finally, US 2004/0230085 A1 discloses a process for treating a hydrocarbon component of biological origin by hydrodeoxygenation followed by isomerization. Notably, U.S. Pat. Nos. 8,865,953 and 8,865,954 disclose processes of making diesel boiling range fuel from renewable sources. The processes in these two patents have several major shortcomings: first they relied on the use of purified hydrogen, which are expensive and energy consuming to produce; second, these processes used excessive amount of hydrogen, partly due to converting the glycerol component in the triglycerides into propane; third, the hydrogen used in these processes are often derived from natural gas, which could be a fossil fuel and a non-renewable source. The use of a large amount of fossil fuel in the production process of a renewable fuel undermines the appeal of the renewable fuel.
The applicant has developed a process which comprises one or more steps to hydrogenate and deoxygenate (via catalytic decarboxylation, decarbonylation and/or hydrodeoxygenation) the feedstock without the use of purified hydrogen. A steam fuel reformer is used to produce a mixture comprising hydrogen and carbon monoxide. Optionally this mixture can go through a water gas shift reactor to convert at least a portion of the carbon monoxide in the mixture to carbon dioxide. The purification of hydrogen by removal of carbon oxides is itself an expensive and energy intensive process. The current invention removes this purification step, and made the water gas shift reactor step optional, thus reducing the costs of hydrogen and the carbon emission associated with the production of hydrogen. In order to further reduce the consumption of hydrogen, the renewable feedstock may optionally go through a hydrolysis process before the hydrogenation process. Glycerol is separated from the sweet water stream at the bottom of the hydrolysis reactor. This prevents the conversion of the glycerol to propane, which is a notable shortcoming of U.S. Pat. Nos. 8,865,953 and 8,865,954. As a result, glycerol is produced instead of propane, and large amount of hydrogen is saved. In order to further reduce hydrogen consumption, the decarboxylation reaction is induced by limiting the amount of hydrogen available after the hydrogenation process and/or by increasing water vapor partial pressure in the deoxygenation process.
Sulfur containing components may be naturally present in the feedstock or may be added to the feedstock or the reaction mixture for various purposes. Carbon dioxide is generated in the reaction zone and need to be at least partially removed from the deoxygenation reactor effluent prior to recycling any excess hydrogen back to the reaction zone. The effluent from the reaction zone is separated into at least a vapor portion and a liquid portion through, for example, cooling and separating. At least some of the liquid portion may be recycled to the reaction zone. The vapor portion is treated using an amine absorber solution to remove at least the carbon dioxide and optionally the sulfur component such as hydrogen sulfide so that the remaining hydrogen can be recycled back to the reaction zone. The separated carbon dioxide and the separated hydrogen sulfide may be used for other purposes. Optionally, a selective separation unit such as a hot high pressure hydrogen stripper may be employed to selectively separate the majority of the hydrocarbon liquid portion from the vapor portion of the effluent and some of this hot, high pressure hydrocarbon liquid portion may be recycled to the reactor. The vapor portion is then cooled to separate any water.
A process for producing a diesel fuel or jet fuel product (also referred to as hydrocarbon fuel, including diesel fuel and/or jet fuel) from a renewable feedstock wherein the process comprises generating a hydrogen-rich mixture in a steam fuel reformer, treating the feedstock in a catalytic reaction zone by hydrogenating and deoxygenating the feedstock at reaction conditions to provide a reaction product comprising a hydrocarbon fraction comprising paraffins and a gaseous fraction comprising at least carbon dioxide and hydrogen. At least one sulfur containing component may be present in the reaction mixture. The sulfur containing component may be present as a naturally occurring contaminant or may be intentionally added to the feedstock or reaction mixture. Many sulfur containing components will form hydrogen sulfide in the reactor. The renewable feedstock may go through an optional hydrolysis process to produce a free fatty acid (FFA) stream and glycerol. The glycerol is separated out and the FFA stream is sent to the reaction zone (also called the deoxygenation reactor) where hydrogenation and deoxygenation occur. It takes 3 molecules of hydrogen to convert a molecule of glycerol to a molecule of propane in a hydrodeoxygenation reaction. The removal of glycerol from the reaction zone therefore reduces the consumption of hydrogen significantly.
The carbon dioxide generated in the reaction zone and any excess hydrogen are selectively removed from the desired reaction product as a vapor stream using, for example, (1) a cooling and separating process or (2) an integrated hot high pressure hydrogen stripper using a hydrogen-rich mixture as the stripping gas followed by the cooling and separating process. The carbon dioxide is then separated from the gaseous effluent stream using at least one selective amine absorber solution. The hydrogen sulfide may be removed from the gaseous effluent stream using the amine absorber solution, or the amine absorber solution may be specially chosen to allow the hydrogen sulfide to recycle with the hydrogen to the reactor. In one embodiment, more than one amine absorber solution may be used in a flexible absorber.
As stated, the present invention relates to a process for producing a hydrocarbon stream useful as diesel boiling range fuel or jet fuel from renewable feedstocks such as renewable feedstocks originating from plants or animals. Some of these feedstocks are known as biorenewable fats and oils. The term renewable feedstock is meant to include feedstocks other than those obtained from crude oil. The renewable feedstocks that can be used in the present invention include any of those which comprise glycerides and/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 include, but are not limited to, canola oil, corn oil, soy oils, 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, jatropha oil, tallow, yellow and brown greases, lard, train oil, fats in milk, fish oil, algal oil, sewage sludge, and the like. Additional examples of renewable feedstocks include non-edible vegetable oils from the group comprising Jatropha curcas (Ratanjoy, Wild Castor, Jangli Erandi), Madhuca indica (Mohuwa), Pongamia pinnata (Karanji Honge), and Azadiracta indicia (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 with a majority of the fats and oils containing high concentrations of fatty acids with 16 and 18 carbon atoms. Mixtures or co-feeds of renewable feedstocks and hydrocarbons from petroleum crude oil may also be used as the feedstock. Other feedstock components which 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, liquids 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. In some applications, an advantage of using a co-feed component is the transformation of what may have been considered to be a waste product from a petroleum based or other process into a valuable co-feed component to the current process.
Renewable feedstocks that can be used in the present invention may contain a variety of impurities. For example, tall oil is a byproduct of the wood processing industry and tall oil contains esters and rosin acids in addition to FFAs. Rosin acids are cyclic carboxylic acids. The renewable feedstocks may also contain contaminants such as alkali metals, e.g. sodium and potassium, phosphorous as well as solids, water and detergents. An optional first step is to remove some or all of these contaminants. One possible pretreatment step involves contacting the renewable feedstock with an ion-exchange resin in a pretreatment zone at pretreatment conditions. The ion-exchange resin is an acidic ion exchange resin such as Amberlyst™-15 and can be used as a bed in a reactor through which the feedstock is flowed through, either upflow or downflow. The conditions at which the reactor is operated are well known in the art.
Another possible means for removing contaminants is a mild acid wash. This is carried out by contacting the feedstock with an acid such as sulfuric, nitric or hydrochloric acid in a reactor. The acid and feedstock can be contacted either in a batch or continuous process. Contacting is done with a dilute acid solution usually at ambient temperature and atmospheric pressure. If the contacting is done in a continuous manner, it is usually done in a counter current manner. Yet another possible means of removing metal contaminants from the feedstock is through the use of guard beds some of which are well known in the art. These can include alumina guard beds either with or without demetallation catalysts such as nickel or cobalt. Filtration and solvent extraction techniques are other choices which may be employed. Hydroprocessing such as that described in U.S. application Ser. No. 11/770,826, incorporated by reference, is another pretreatment technique which may be employed.
Hydrogen is typically produced using a methane steam reformer (SMR). According to “Refinery Process Modeling” 1st edition by Gerald Kaes, Kaes Enterprises Inc (2000), natural gas is reacted with water vapor to produce hydrogen in multiple steps, comprising at least the steps of: 1) reforming, typically performed at 20-30 atm and 800-880° C.; 2) water gas shift reaction, possibly in two steps, a high temperature shift reaction at 345-370° C. and a low temperature shift reaction at about 230° C.; and 3) hydrogen purification, using Pressure Swing Adsorption (PSA), membrane or amine absorption. The reforming reaction is endothermic, requiring input of external heat. The heat is typically provided by burning fossil fuel such as natural gas. Hydrogen is an expensive resource that does not exist naturally, and therefore it is highly desirable to minimize the use of hydrogen and reduce the cost of hydrogen.
The reaction that occurs in the reforming step of a SMR is:
CH4+H2O→CO+3H2 (1)
The reforming reaction product mixture in reality may contain hydrogen, carbon monoxide, unreacted water, and some carbon dioxide. If the concentrations of water and carbon dioxide are low or excluded, the product mixture is about 75% hydrogen and 25% carbon monoxide. Therefore, the product mixture from the reforming step is referred to as a hydrogen-rich mixture.
The reaction that occurs in the water shift reaction in the hydrogen production process is:
CO+H2O→CO2+H2 (2)
Combining reaction (1) and reaction (2) yields an overall reaction:
CH4+2H2O→CO2+4H2 (3)
The product mixture after the water shift reaction is mostly hydrogen and carbon dioxide. In reality the product mixture may contain some unreacted water and carbon monoxide. But if the concentrations of water and carbon monoxide are low or excluded, the product mixture is about 80% hydrogen and 20% carbon dioxide. Again the product mixture is referred to as a hydrogen-rich mixture. The purification of hydrogen itself is expensive in terms of capital costs and operating costs. It also consumes energy and release the greenhouse gas carbon dioxide. The current invention removes the purification step. It also makes the water gas shift reaction optional. When the water gas shift reaction is skipped, a hydrogen-rich mixture comprising hydrogen and carbon monoxide is used in lieu of purified hydrogen in the production of biorenewable diesel fuel or jet fuel. Even though the optional water gas shift reactor is removed, the water gas shift reaction still need to occur; it now occurs in the reaction zone (in the deoxygenation reactor). The deoxygenation reactor can assume multiple functions include hydrogenation, deoxygenation, and water gas shift reaction. This again saves capital costs.
The renewable feedstock is flowed to a reaction zone comprising one or more catalyst beds in one or more reactors. The term “feedstock” is meant to include feedstocks that have not been treated to remove contaminants as well as those feedstocks purified in a pretreatment zone. In the reaction zone, the feedstock is contacted with a hydrogenation or hydrotreating catalyst in the presence of hydrogen at hydrogenation conditions to hydrogenate reactive component such as the olefinic or unsaturated portions of the aliphatic chains of a triglyceride or FFA molecule. Hydrogenation and hydrotreating catalysts are any of those well known in the art such as nickel or nickel/molybdenum dispersed on a high surface area support. Other hydrogenation catalysts include one or more noble metal catalytic elements dispersed on a high surface area support. Non-limiting examples of noble metals include Pt and/or Pd dispersed on gamma-alumina. Hydrogenation conditions include a temperature of about 40° C. to about 400° C. and a pressure of about 689 kPa absolute (100 psia) to about 13,790 kPa absolute (2000 psia). In another embodiment the hydrogenation conditions include a temperature of about 200° C. to about 300° C. and a pressure of about 1379 kPa absolute (200 psia) to about 4826 kPa absolute (700 psia). Other operating conditions for the hydrogenation zone are well known in the art.
The hydrogenation and hydrotreating catalysts enumerated above are also capable of catalyzing decarbonylation, decarboxylation and/or hydrodeoxygenation of the feedstock to remove oxygen. Decarbonylation, decarboxylation, hydrodeoxygenation and hydrogenation are herein collectively referred to as deoxygenation reactions. Decarbonylation, decarboxylation, and hydrodeoxygenation conditions include a relatively low pressure of about 1379 kPa (200 psia) to about 6895 kPa (1000 psia), a temperature of about 200° C. to about 400° C. and a liquid hourly space velocity of about 0.5 to about 10 hr−1. In another embodiment the deoxygenation conditions include the same relatively low pressure of about 1379 kPa (200 psia) to about 6895 kPa (1000 psia), a temperature of about 288° C. to about 345° C. and a liquid hourly space velocity of about 1 to about 4 hr−1. Since hydrogenation is an exothermic reaction, as the feedstock flows through the catalyst bed the temperature increases and deoxygenation will begin to occur. Thus, it is envisioned and is within the scope of this invention that all the reactions may occur simultaneously in one reactor or in one bed. Alternatively, the conditions can be controlled such that hydrogenation primarily occurs in one bed and deoxygenation occurs in a second bed. Of course if only one bed is used, then hydrogenation occurs primarily at the front end of the bed, while deoxygenation occur mainly in the back end of the bed. Finally, desired hydrogenation can be carried out in one reactor, while decarboxylation and/or hydrodeoxygenation can be carried out in a separate reactor. This configuration is desirable since the hydrogen level between the two reactors can be monitored and controlled to a relatively low level, to induce the decarboxylation reaction and suppress the hydrodeoxygenation reaction in order to reduce hydrogen consumption.
Hydrolysis of triglycerides is commonly carried out to produce free fatty acids and a byproduct glycerol. The process typically used today is a variation of the so-called Colgate-Emery process, according Barnebey, H. & Brown, A. C. (1948), Continuous Fat Splitting Plants Using the Colgate-Emery Process. J. Am. Oil Chem. Soc. Vol. 25, pp. 95-99, ISSN 0003-021X. The Colgate-Emery process is typically carried out at up to pressure of 5 MPa (725 psia) and temperature of 260° C. U.S. Pat. No. 4,218,386 disclosed a more recent process for hydrolyzing triglycerides.
The renewable feedstock may go through an optional hydrolysis step to produce free fatty acids (FFA) and glycerol. The glycerol is separated out and FFA is sent to the deoxygenation reactor where hydrogenation and deoxygenation occur. It takes 3 molecules of hydrogen to convert a molecule of glycerol to propane in a hydrodeoxygenation reaction. The removal of glycerol from the reaction zone reduces the consumption of hydrogen significantly. In order to reduce the hydrogen consumption, decarboxylation is preferred over hydrodeoxygenation. Since it requires hydrogen to carry out the hydrodeoxygenation reaction, it is possible to control and limit the amount of hydrogen available after the hydrogenation process and before the deoxygenation process. Since water vapor is a product of hydrodeoxygenation, it is also possible to suppress the hydrodeoxygenation reaction in favor of the decarboxylation reaction by increasing the partial pressure of water vapor in the de-oxygenation reaction zone.
Sulfur containing components are often present in the reaction mixture. Such components may be present in the feedstock naturally, or may be added to the feedstock or the reaction zone. Sulfur-containing components may be organic, inorganic, natural, or synthetic. A single sulfur-containing component may be present or more than one may be present. The sulfur containing component may be present in an amount ranging from about 1 ppm to about 5 mass %. Many sulfur containing components are converted to hydrogen sulfide in the reaction zone. For ease of understanding, the description below will use the term hydrogen sulfide as the primary example of a sulfur containing component, but that is not meant to limit the scope of the claims in any way.
The reaction product from the deoxygenation reactions comprises both a liquid portion and a gaseous portion. The liquid portion comprises a hydrocarbon fraction which is primarily paraffins and having a large concentration of paraffins in the range of about 9 to about 18 carbon atoms. Different feedstocks will result in different distributions of paraffins. The gaseous portion comprises hydrogen, carbon dioxide, carbon monoxide, water vapor, and perhaps sulfur components such as hydrogen sulfide or phosphorous component such as phosphine. For the case where there is no isomerization catalyst in the reaction zone, most of the hydrocarbons will be normal paraffins. In this case, the hydrogenation or hydrotreating/deoxygenation catalyst may catalyze a slight amount of isomerization but it is expected that no more than about 5 or about 10 mass % of the normal paraffins would be isomerized to branched paraffins. The cold flow properties of the diesel boiling range fuel or jet fuel depend on the relative amounts of normal and branched paraffin in the product. In warmer climate regions, poor cold flow properties are not a great concern. In colder climate regions, improvements to cold flow properties are needed and at least some of the normal paraffins are isomerized to branched paraffins. By optimizing the isomerization requirement where appropriate due to the climate, a substantial cost savings in both capital costs and operating costs can be achieved.
The effluent from the reaction zone is conducted to a selective separation zone comprising, for example, a heat exchanger and a product separator and optionally an air or water cooler. After cooling, a vapor stream containing the hydrogen, hydrogen sulfide, carbon monoxide, and carbon dioxide is readily separated from the liquid phase containing the normal paraffins having from about 8 to about 24 carbon atoms in the product separator. Suitable operating conditions of the separator include, for example, a temperature of about 20 to 80° C. or 40 to 50° C. and a pressure of about 2758 kPa absolute (400 psia) to about 68985 kPa absolute (1000 psia) with a specific embodiment at 3850 kPa absolute (560 psia). This selective separation zone is operated at essentially the same pressure as the reaction zone. By “essentially” it is meant that the operating pressure of the selective separation zone is within about 1034 kPa absolute (150 psia) of the operating pressure of the reaction zone. For example, the selective separation zone is no more than 1034 kPa absolute (150 psia) less than that of the reaction zone. The vapor stream and the liquid stream are both removed from the product separator. A portion of the liquid stream may be recycled to the reaction zone, at the feed location or at one or more intermediate locations. A water byproduct stream is also removed. The liquid stream may be recovered or may be routed to a product recovery column to separate the light ends from the diesel and naphtha products.
Optionally, the effluent from the reaction zone is conducted to a hot high pressure hydrogen stripper before at least a portion of the effluent is cooled and conducted to the cold product separator. One benefit of this embodiment is that a liquid stream of paraffins is generated at or near to the temperature and pressure of the reaction zone, and a portion of that stream may be recycled to the reaction zone with minimal pumping energy and minimal additional heating. Saving the utilities of pumping and reheating can significantly reduce the cost of the overall process and if the recycle stream is large enough would more than offset the additional capital cost of the hot high pressure hydrogen stripper. Likewise the net liquid going to the product recovery column needs less heating to separate light byproducts. In addition, the separation in the cold product separator becomes more efficient since the phase separation does not include the densest hydrocarbons having from about 8 to about 24 or more carbon atoms. Furthermore, any unreacted glycerides or free fatty acids present in the reactor effluent during a unit start-up or unit upset are selectively removed in the hot separator liquid and do not come into contact with a condensed water phase where they could contaminate the byproduct water.
The reaction zone effluent enters the hot high pressure hydrogen stripper and the water and normally gaseous components, are carried with the hydrogen-rich stripping gas and separated into an overhead stream. By using a hydrogen-rich mixture stream from the steam fuel reformer as the stripping gas (in lieu of purified hydrogen), water, carbon monoxide, carbon dioxide, and any ammonia or hydrogen sulfide are selectively separated from the hydrocarbon liquid product in the hot high pressure hydrogen stripper. The remainder of the deoxygenation effluent stream is removed as hot high pressure hydrogen stripper bottoms and contains the liquid hydrocarbon fraction having components such as normal hydrocarbons having from about 8 to about 24 carbon atoms. A portion of this liquid hydrocarbon fraction in hot high pressure hydrogen stripper bottoms may be used as the hydrocarbon recycle described below, and the stripper bottoms are already at or near the operating conditions of the reaction zone thereby saving the costs involved with pumping or heating of the recycle portion. The stripper bottoms are conducted to a product recovery column.
The temperature of the hot high pressure hydrogen stripper may be controlled in a limited range to achieve the desired separation and the pressure may be maintained at approximately the same pressure as the reaction zone to minimize both investment and operating costs. The hot high pressure hydrogen stripper may be operated at conditions ranging from a pressure of about 689 kPa absolute (100 psia) to about 13,790 kPa absolute (2000 psia), and a temperature of about 40° C. to about 350° C. In another embodiment the hot high pressure hydrogen stripper may be operated at conditions ranging from a pressure of about 1379 kPa absolute (200 psia) to about 4826 kPa absolute (700 psia), or about 2413 kPa absolute (350 psia) to about 4882 kPa absolute (650 psia), and a temperature of about 50° C. to about 350° C. The hot high pressure hydrogen stripper may be operated at essentially the same pressure as the reaction zone. By “essentially” it is meant that the operating pressure of the high pressure hydrogen stripper is within about 1034 kPa absolute (150 psia) of the operating pressure of the reaction zone. For example, the pressure of the hot high pressure hydrogen stripper separation zone is no more than 1034 kPa absolute (150 psia) less than that of the reaction zone. Also, the stream entering the hot high pressure hydrogen stripper may be heat exchanged to reduce the temperature before entering the hot high pressure hydrogen stripper. In this way the optimum temperature value needed to achieve the selective separation is obtained prior to entering the hot high pressure hydrogen stripper.
One purpose of the hot high pressure hydrogen stripper is to separate the gaseous portion of the effluent from the liquid portion of the effluent. As hydrogen is an expensive resource, to conserve costs, the separated hydrogen-rich gas mixture is ultimately recycled to the deoxygenation reactor after carbon dioxide and at least some hydrogen sulfide are removed. Since carbon monoxide is a regulated pollutant and considered a fuel, releasing it to the atmosphere is not desirable. Therefore, carbon monoxide (if any exist in the gaseous effluent) should be recycled with the hydrogen-rich mixture to the reaction zone. Hydrogen is a reactant in at least one of the deoxygenation reactions, and to be effective, a sufficient quantity of hydrogen must be in solution to most effectively take part in the catalytic reaction. Past processes have operated at high pressures in order to achieve a desired amount of hydrogen in solution that is readily available for reaction. However, higher pressure operations are more costly to build and to operate as compared to their lower pressure counterparts. One advantage of the present invention is the ability to operate in a pressure range of about 1379 kPa absolute (200 psia) to about 4826 kPa absolute (700 psia) which is lower than that found in other previous operations. In another embodiment the operating pressure is in the range of about 2413 kPa absolute (350 psia) to about 4481 kPa absolute (650 psia), and in yet another embodiment operating pressure is in the range of about 2758 kPa absolute (400 psia) to about 4137 kPa absolute (600 psia). Furthermore, the rate of reaction is increased resulting in a greater amount of throughput of material through the reactor in a given period of time.
In one embodiment, the desired amount of hydrogen is kept in solution at lower pressures by employing a large recycle of hydrocarbon. Other processes have employed hydrocarbon recycle in order to control the temperature in the reaction zones since the reactions are exothermic reactions. However, the range of recycle to feedstock ratios used herein is determined not on temperature control requirements, but instead, based upon feedstock composition and hydrogen solubility requirements. Hydrogen has a greater solubility in the hydrocarbon product than it does in the feedstock. By utilizing a large hydrocarbon recycle the solubility of hydrogen in the liquid phase in the reaction zone is greatly increased and higher pressures are not needed to increase the amount of hydrogen in solution. In one embodiment of the invention, the volume ratio of hydrocarbon recycle to feedstock is from about 2:1 to about 8:1. In another embodiment the ratio is in the range of about 3:1 to about 6:1 in yet another embodiment the ratio is in the range of about 4:1 to about 5:1, and in still another embodiment the ratio is in the range of about 2:1 to about 6:1.
The gaseous portion of the reaction zone effluent in the overhead from the hot high pressure hydrogen stripper is cooled, by techniques such as heat exchange, air cooling, or water cooling and passed to a cold separator where liquid components are separated from the gaseous components by phase separation. Suitable operating conditions of the cold separator include, for example, a temperature of about 20 to 80° C. or 40 to 50° C. and a relatively low pressure of about 3447 kPa (500 psia) to about 6895 kPa (1000 psia), with one embodiment at 3850 kPa absolute (560 psia). A water byproduct stream is also separated. The gaseous component stream from the cold separator comprises hydrogen, carbon monoxide, carbon dioxide, and hydrogen sulfide while the liquid component stream from the cold separator comprises naphtha and LPG (if the optional hydrolysis step is skipped). Again, this separation may be operated at essentially the same pressure as the reaction zone. By “essentially” it is meant that the operating pressure of the cold separator is within about 1034 kPa absolute (150 psia) of the operating pressure of the reaction zone. For example, the pressure of the separator is no more than 1034 kPa absolute (150 psia) less than that of the reaction zone.
Either the hot high pressure hydrogen stripper bottoms or the liquid component from the cold product separator in the embodiment with no hot high pressure hydrogen stripper may be recovered as diesel or jet fuel product. However, the liquid component from the product separator and the hot high pressure hydrogen stripper bottoms, if present, collectively contain the hydrocarbons useful as diesel boiling range fuel or jet fuel as well as smaller amounts of naphtha and LPG (if any) and may be further purified in a product recovery column. The product recovery column fractionates lower boiling components and dissolved gases from the diesel product containing C8 to C24 paraffins. Suitable operating conditions of the product recovery column include a temperature of from about 20 to about 200° C. at the overhead and a pressure from about 0 to about 1379 kPa absolute (0 to 200 psia).
The naphtha and LPG (if any) stream may be further separated in a debutanizer or depropanizer in order to separate the LPG into an overhead stream, leaving the naphtha in a bottoms stream. In the embodiment with a hydrolysis step included before the hydrogenation and deoxygenation, the amount of LPG produced in the deoxygenation reactor may be very little, if any. Suitable operating conditions of this unit include a temperature of from about 20 to about 200° C. at the overhead and a pressure from about 0 to about 2758 kPa absolute (0 to 400 psia). The glycerol from the optional hydrolysis step and/or naphtha from the hydrocarbon product stream may be used as either hydrogen donor fuel or the fuel used in a combustion process to generate the heat for the steam fuel reformer. A portion of the feedstock may be used as the fuel in a combustion process to generate the heat for the steam fuel reformer. By utilizing glycerol and other hydrocarbons in the steam fuel reforming process, it is possible to produce a renewable diesel fuel or jet fuel without use of any fossil-fuel including natural gas. Of course, when natural gas is widely available and affordable, it may still be used in the steam fuel reformer to increase the overall competitiveness of the process of producing renewable fuels. In fact natural gas is often difficult to transport to other locations and sometimes wasted by flaring. The current invention may serve as a convenient way to make use of the abundant natural gas.
The gaseous component separated in the product separator of either embodiment above comprises mostly hydrogen and the carbon dioxide from the decarboxylation reaction. Other components such as carbon monoxide, and hydrogen sulfide or other sulfur containing component may be present as well. It is desirable to recycle the hydrogen and carbon monoxide (if any) to the reaction zone, but if the carbon dioxide was not removed, its concentration would quickly build up and effect the operation of the reaction zone. Usually, carbon dioxide would be removed from the hydrogen by means well known in the art such as absorption, along with hydrogen sulfide, using an amine, reaction with a hot carbonate solution, pressure swing adsorption, etc. and if desired, essentially pure carbon dioxide could be recovered by regenerating the spent absorption media. However, the separation of carbon dioxide from hydrogen is complicated by the sulfur containing component such as hydrogen sulfide which is present to maintain the sulfided state of the deoxygenation catalyst or to control the relative amounts of the decarboxylation reaction and the hydrogenation reaction that are both occurring in the deoxygenation zone. Because the hydrogen sulfide serves a useful purpose in the reaction zone, it is desirable to recycle the hydrogen sulfide to the reaction zone as opposed to purchasing additional hydrogen sulfide or sulfur components. In some applications, there may be a need to control the level of hydrogen sulfide being recycled which may require removing substantially all the hydrogen sulfide in order to control the amount of separated hydrogen sulfide that is recycled to the reaction zone. Therefore, the techniques for removing the carbon dioxide also need to provide the sulfur management in the process.
In one embodiment of the invention an amine absorber is used to selectively remove carbon dioxide while allowing hydrogen, and hydrogen sulfide to pass to recycle. In this embodiment the gaseous stream from the cold product separator is routed through an amine absorber containing an aqueous solution of a polyoxypropylene triamine having the formula:
Where R′ represents a methylene group and R″ represents hydrogen or methyl or ethyl and wherein the sum of X+Y=Z is a positive integer having a value of from about 4 to about 6. These amines are fully described in U.S. Pat. No. 4,710,362 which is hereby incorporated by reference in its entirety. The amine is in an aqueous solution containing about 35 to about 55 wt. % of the polyoxypropylene triamine, and the absorption in the absorber may be conducted at about 20° C. to about 50° C.
In another embodiment, two amine absorbers are employed. The first amine scrubber removes both carbon dioxide and hydrogen sulfide allowing hydrogen to pass to recycle. The amine chosen to be employed in first amine absorber is capable of removing at least both the components of interest, carbon dioxide and the sulfur components such as hydrogen sulfide. Suitable amines are available from DOW and from BASF, and in one embodiment the amines are a promoted or activated methyldiethanolamine (MDEA). The promoter may be piperazine, and the promoted amine may be used as an aqueous solution. See U.S. Pat. No. 6,337,059, hereby incorporated by reference in its entirety. Suitable amines for the first amine absorber from DOW include the UCARSOL™ AP series solvents such as AP802, AP804, AP806, AP810 and AP814. The carbon dioxide and hydrogen sulfide are absorbed by the amine while the hydrogen passes through first amine absorber to be recycled to the reaction zone. The amine is regenerated and the carbon dioxide and hydrogen sulfide are released and removed. The regenerated amine may be recycled and reused. The released carbon dioxide and hydrogen sulfide are passed through a second amine absorber which contains an amine selective to hydrogen sulfide, but not selective to carbon dioxide. Again, suitable amines are available from DOW and from BASF, and in one embodiment the amines are a promoted or activated MDEA. Suitable amines for the second amine absorber zone from DOW include the UCARSOL™ HS series solvents such as HS101, HS102, HS103, HS104, HS115. Therefore the carbon dioxide passes through second amine absorber and is available for use elsewhere. The amine may be regenerated which releases the hydrogen sulfide to be recycled. A portion of the hydrogen sulfide may be sent to a Claus plant. Regenerated amine is then recycled and reused. The hydrogen sulfide recycle to the reaction zone may be controlled so that the appropriate amount of sulfur is maintained in the reaction zone. Conditions for the first scrubber zone includes a temperature in the range of about 30 to about 60° C. At least the first absorber is operated at essentially the same pressure as the reaction zone. By “essentially” it is meant that the operating pressure of the absorber is within about 1034 kPa absolute (150 psia) of the operating pressure of the reaction zone. For example, the pressure of the absorber is no more than about 1034 kPa absolute (150 psia) less than that of the reaction zone. Also, at least the first absorber is operated at a temperature that is at least about 1° C. higher than that of the separator. Keeping at least the first absorber warmer than the separator operates to maintain any light hydrocarbons in the vapor phase and prevents the light hydrocarbons from condensing into the absorber solvent. Conditions for the second amine solution absorber zone may include from about 20 to about 60° C. and a pressure in the range of about 138 kPa (20 psia) to about 241 kPa (35 psia).
The gaseous component stream from the cold product separator has a total volume that is much greater than the combined volume of carbon dioxide and hydrogen sulfide. Typically, the amount of hydrogen sulfide in vapor stream 236, 336 or 436 ranges from about 0.01 to about 2 volume-%. In the configurations shown in the figures, the first amine absorber zone is sized to accommodate the flow of the entire vapor stream from the cold product separator. However, the second amine absorber zone is greatly reduced in size as compared to the first since the flow of material to the second amine absorber zone is only a fraction of vapor stream from the cold product separator. The reduction in the size of the second amine absorber zone allows for reduced capital and operating costs.
Other separation systems are possible, such as adsorbents and treating processes. However, the amine absorber systems of the present invention have several advantages with cost being a primary advantage. Amine absorber systems are less costly than molecular sieve adsorbents or treating processes, and the amine systems minimize the amount of hydrogen lost to the acid gas containing stream(s).
The hydrogen stream remaining after the removal of the carbon dioxide may contain carbon monoxide, and may be recycled to the reaction zone. The hydrogen stream may contain the hydrogen sulfide being recycled to the reaction zone, or the separated hydrogen sulfide may be recycled independently such as in controlled amounts. The hydrogen recycle stream may be introduced to the inlet of the reaction zone and/or to any subsequent beds/reactors. To facilitate the water gas shift reaction in the reaction zone (for the conversion of carbon monoxide to carbon dioxide and hydrogen), steam may be injected into the reaction zone in various locations, such as upstream of the reaction zone, or in between stages of the reaction zone.
The following embodiments are presented in illustration of this invention and are not intended as an undue limitation on the generally broad scope of the invention as set forth in the claims. First the process is described in general as with reference to
Turning to
Hydrogen donor fuel 12 and hot steam 14 enter the steam-fuel reformer 10, and under high temperature conditions (between 700 and 1100° C., or more specially between 800 and 880° C.) produces a gas mixture stream 16 comprising hydrogen and carbon monoxide. The most commonly used hydrogen donor fuel is natural gas. Since the main component of natural gas is methane, reactor 10 is often referred to as Steam Methane Reformer (SMR) in the existing arts. But the hydrogen donor fuel is not limited to natural gas or methane, and therefore the general name “steam fuel reformer” is used in this invention. For example, glycerol or naphtha can be used as the hydrogen donor fuel as well. Since steam fuel reforming is an endothermic process, steam fuel reformer 10 requires external source of heat, which is typically supplied by a combustion process, which is well known in the arts and not shown. This combustion process could use natural gas or any other commonly available fuel, such as #2 oil and propane. However, it is preferable that the fuel used in this combustion process uses a renewable source as well, if it is economically feasible to do so, and the renewable source is readily available. For example, glycerol, naphtha or even some of the glycerides could be burned in the combustion process to generate the heat needed for reactor 10. If the donor fuel 12 is glycerol and naphtha from certain processes of the current invention, and the fuel used for generating the heat needed for reformer 10 is also from a renewable source, it is possible to produce a renewable diesel or jet fuel that is completely renewable, without the use of fossil fuel in the process.
Hydrogen-rich mixture stream 16 can go through an optional water gas shift reactor 20, under a lower temperature than the steam fuel reformer 10 and in the presence of steam, produce a stream 22 comprising mostly hydrogen and carbon dioxide. An optional bypass line 18 can be used to bypass reactor 20 if higher level of carbon monoxide is acceptable or is desirable. Hydrogen-rich mixture stream 22 and optionally recycle hydrogen is added to optional first selective separation zone 108.
Carbon oxides, hydrogen sulfide, and water vapor are removed with hydrogen-rich mixture stream 22 in optional first selective separation zone overhead 114 and separated deoxygenated liquid product are removed in optional first selective separation zone bottoms 118. Both streams 114 and 118 are passed to product recovery zone 120. Product recovery zone 120 comprises at least a cooler, a cold product separator, and a product recovery column. Carbon oxides, hydrogen sulfide, and hydrogen stream 122, light ends stream 124, water byproduct stream 128, and paraffin-rich product 119 are all removed from product recovery zone 120. Paraffin-rich product 119 may be collected for use as diesel boiling range fuel or jet fuel. Stream 122 is directed to second selective separation zone 130 which contains one or more selective amine absorbers. At least carbon dioxide is removed from stream 122 via line 132. Optionally hydrogen sulfide may be removed as well. Hydrogen recycle stream 126 is removed from second selective separation zone and recycled to the reaction zone 104.
Turning to
Hydrogen donor fuel 12 and hot steam 14 enter the steam-fuel reformer 10, and under high temperature conditions (between 700 and 1100° C., or more specially between 800 and 880° C.) produces a gas mixture stream 16 comprising hydrogen and carbon monoxide.
Hydrogen-rich mixture stream 16 can go through an optional water gas shift reactor 20, under a lower temperature than the steam fuel reformer 10 and in the presence of steam, produce a stream 22 comprising mostly hydrogen and carbon dioxide. An optional bypass line 18 can be used to bypass reactor 20 if higher level of carbon monoxide is acceptable or is desirable. Hydrogen-rich mixture stream 22 and optionally recycle hydrogen is added to optional first selective separation zone 108.
Carbon oxides, hydrogen sulfide, and water vapor are removed with hydrogen-rich mixture stream 22 in optional first selective separation zone overhead 114 and separated deoxygenated liquid product are removed in optional first selective separation zone bottoms 115. Stream 115, together with a portion of hydrogen-rich mixture stream 22 through line 24, is routed to an isomerization reactor 116 to convert a portion of n-parafins to iso-parafins. A portion of the recycled stream 126 can also be routed to the isomerization reactor 116 through line 126a. The product stream 118 from the isomerization reactor 116 is combined with overhead stream 114, then passed to product recovery zone 120. Product recovery zone 120 comprises at least a cooler, a cold product separator, and a product recovery column. Carbon oxides, hydrogen sulfide, and hydrogen stream 122, light ends stream 124, water byproduct stream 128, and paraffin-rich product 119 are all removed from product recovery zone 120. Paraffin-rich product 119 may be collected for use as diesel boiling range fuel or jet fuel. Stream 122 is directed to second selective separation zone 130 which contains one or more selective amine absorbers. At least carbon dioxide is removed from stream 122 via line 132. Optionally hydrogen sulfide may be removed as well. Hydrogen recycle stream 126 is removed from second selective separation zone and recycled to the reaction zone 104.
Turning to
Renewable feedstock stream 202 may pass through an optional feed surge drum. The renewable feedstock 202 may be glycerides and free fatty acids if an optional hydrolysis step is not included, but it can be mostly free fatty acids if the optional hydrolysis step is included (not shown in
It should be noted the steam fuel reformer 10 supplies the hydrogen-rich mixture 18 as the source of hydrogen to the deoxygenation zone in two ways, both directly and indirectly, one directly supplying to the reaction zone 204 through stream 18, and the other indirectly to the reaction zone 204 by first going through other processes before going to the reaction zone through recycle stream 276.
Deoxygenation reactor 204 may contain multiple beds shown in
Deoxygenation reactor effluent stream 206, after one or more optional heat exchanges, is directed to air cooler 232 and then introduced into product separator 234. In product separator 234 the gaseous portion of the stream comprising hydrogen, carbon monoxide, hydrogen sulfide and carbon dioxide are phase separated and removed in stream 236 while the liquid hydrocarbon portion of the stream is removed in stream 238. A portion of the liquid hydrocarbon stream 238a is recycled to the reaction zone 204. A water byproduct stream 240 may also be removed from product separator 234. Stream 238 is introduced to product recovery column 242 where components having higher relative volatilities are separated into stream 244 with the remainder, the diesel or jet fuel range components, being withdrawn from product recovery column 242 in line 246. Stream 244 is optionally introduced into a fractionator which operates to separate LPG (if any) into an overhead leaving a naphtha bottoms stream (not shown).
The vapor stream 236 from product separator 234 contains the gaseous portion of the reaction zone effluent which comprises at least hydrogen, carbon monoxide, hydrogen sulfide and carbon dioxide and is directed to a system of at least one amine absorber and regenerator 256 to separate carbon dioxide and optionally hydrogen sulfide from the vapor stream. Because of the cost of hydrogen, it is desirable to recycle the hydrogen to deoxygenation reactor 204, but it is not desirable to circulate the carbon dioxide or an excess of sulfur containing components. In one embodiment, vapor stream 236 is passed through a system of one amine absorber 256, also called a scrubber. The amine chosen to be employed in the single amine absorber 256 is capable of selectively removing carbon dioxide while allowing hydrogen and hydrogen sulfide to pass through the absorber. Suitable amines for use in are described in U.S. Pat. No. 4,710,362. The amine absorber may be operated at from about 20 to about 60° C. and a pressure in the range of about 3447 kPa (500 psia) to about 6895 kPa (1000 psia).
In another embodiment, to separate both the sulfur containing components and the carbon dioxide from the hydrogen, vapor stream 236 is passed through a system of at least two amine absorbers 256 and 258. The amine employed in amine absorber 256 is capable of selectively removing at least both the components of interest, carbon dioxide and the sulfur components such as hydrogen sulfide. Suitable amines are available from DOW and from BASF, and in one embodiment the amines are a promoted or activated methyldiethanolamine (MDEA). The promoter may be piperazine, and the promoted amine may be used as an aqueous solution. See U.S. Pat. No. 6,337,059, hereby incorporated by reference in its entirety. Suitable amines for the first amine absorber zone from DOW include the UCARSOL™ AP series solvents such as AP802, AP804, AP806, AP810 and AP814. The carbon dioxide and hydrogen sulfide are absorbed by the amine while the hydrogen passes through first amine scrubber zone and into line 276 to be recycled to reaction zone 204. The amine is regenerated and the carbon dioxide and hydrogen sulfide are released and removed in line 262. Within the first amine absorber zone, regenerated amine may be recycled for use again. The released carbon dioxide and hydrogen sulfide in line 262 are passed through optional second amine scrubber zone 258 which contains an amine selective to hydrogen sulfide, but not selective to carbon dioxide. Again, suitable amines are available from DOW and from BASF, and in one embodiment the amines are a promoted or activated MDEA. Suitable amines for the second amine absorber zone from DOW include the UCARSOL™ HS series solvents such as HS101, HS102, HS103, HS104, HS115. Therefore the carbon dioxide passes through second amine scrubber zone 258 and into line 266. The amine may be regenerated which releases the hydrogen sulfide into line 260. At least a portion of the hydrogen sulfide in line 260 may be recycled to the reaction zone 204, possibly in measured controlled amount. Excess hydrogen sulfide may be directed to a Claus plant. Regenerated amine is then reused. Conditions for the first amine solution absorber zone includes from about 30 to about 60° C. and a pressure in the range of about 3447 kPa (500 psia) to about 6895 kPa (1000 psia) and conditions for the second amine solution absorber zone (258) includes from about 20 to about 60° C. and a pressure in the range of about 138 kPa (20 psia) to about 241 kPa (35 psia).
In another embodiment, the amine solution absorber zone 256 may contain the amine solution of U.S. Pat. No. 4,710,362 which selectively separates only the carbon dioxide and allows the hydrogen sulfide to pass with the hydrogen into recycle line 276. In this embodiment, the second amine absorber zone 258 is not necessary.
Another embodiment of the invention employs a hot high pressure hydrogen stripper. Turning to
The feedstock stream 302 is combined with recycle stream 376 to form combined feed stream 320, which is heat exchanged with reactor effluent and then introduced into deoxygenation reactor 304. The heat exchange may occur before or after the recycle is combined with the feed. Deoxygenation reactor 304 may contain multiple beds shown in
Hydrogen donor fuel 12 and hot steam 14 enter the steam-fuel reformer 10, and under high temperature conditions (typically between 700 and 1100° C., and more specifically 800-880° C.) produces a gas mixture stream 16 comprising hydrogen and carbon monoxide. Stream 16 can go through an optional water gas shift reactor 20, under lower temperature conditions than the steam fuel reformer 10, and in the presence of steam, produce a stream 22 comprising hydrogen and carbon dioxide. The water gas shift reaction in reactor 20 is slightly exothermic, converting carbon monoxide and water to hydrogen and carbon dioxide. A typical hydrogen production system comprises three subsystems: the steam methane reformer, the water gas shift reactor, and a hydrogen purification system. The current invention removes the hydrogen purification step, and even made the water gas shift reactor optional, thus reducing costs and energy consumption.
Deoxygenation reactor effluent stream 306 is directed to hot high pressure hydrogen stripper 308. Hydrogen-rich mixture stream 22 is also introduced to hot high pressure hydrogen stripper 308. In hot high pressure hydrogen stripper 308, the gaseous component of deoxygenation reactor effluent 306 is stripped from the liquid component of deoxygenation reactor effluent 306 using hydrogen-rich mixture stream 22 and optional recycle hydrogen (not shown). The gaseous component comprising hydrogen, vaporous water, carbon monoxide, carbon dioxide and possibly some propane, is separated into hot high pressure hydrogen stripper overhead stream 314. The remaining liquid component of deoxygenation reactor effluent 306 comprising primarily normal paraffins having a carbon number from about 8 to about 24 with a cetane number of about 60 to about 100 is removed as hot high pressure hydrogen stripper bottoms 312.
A portion of hot high pressure hydrogen stripper bottoms forms recycle stream 313 and is combined with renewable feedstock combined stream 320. Another portion of recycle stream 313, may be routed directly to deoxygenation reactor 304 and introduced at interstage locations such as between beds 304a and 304b and/or between beds 304b and 304c, in order, for example, to aid in temperature control or enhance soluble hydrogen level in the reactants. The remainder of hot high pressure hydrogen stripper bottoms in stream 312 is routed to product recovery column 342.
Hydrogen stripper overhead stream 314 is air cooled using air cooler 332 and introduced into product separator 334. In product separator 334 the gaseous portion of the stream comprising hydrogen, carbon monoxide, hydrogen sulfide and carbon dioxide are removed in stream 336 while the liquid hydrocarbon portion of the stream is removed in stream 338. A liquid water byproduct stream 340 may also be removed from product separator 334. Stream 338 is introduced to product recovery column 342 where components having higher relative volatilities are separated into stream 344 with the remainder, the diesel range components, being withdrawn from product recovery column 342 in line 346. Stream 344 may be introduced into a fractionator which operates to separate LPG into an overhead and leaving a naphtha bottoms (not shown).
The vapor stream 336 from product separator 334 contains the gaseous portion of the deoxygenation effluent which comprises at least hydrogen, carbon monoxide, hydrogen sulfide and carbon dioxide and is directed to a system of amine absorbers to separate carbon dioxide and hydrogen sulfide from the vapor stream. Because of the cost of hydrogen, it is desirable to recycle the hydrogen to deoxygenation reactor 304, but it is not desirable to circulate the carbon dioxide or an excess of sulfur containing components. In order to separate sulfur containing components and carbon dioxide from the hydrogen, vapor stream 336 is passed through a system of at least two amine absorbers, also called scrubbers, starting with the first amine absorber zone 356. The amine chosen to be employed in first amine scrubber 356 is capable of selectively removing at least both the components of interest, carbon dioxide and the sulfur components such as hydrogen sulfide. Suitable amines are available from DOW and from BASF, and in one embodiment the amines are a promoted or activated methyldiethanolamine (MDEA). The promoter may be piperazine, and the promoted amine may be used as an aqueous solution. See U.S. Pat. No. 6,337,059, hereby incorporated by reference in its entirety. Suitable amines for the first amine absorber zone from DOW include the UCARSOL™ AP series solvents such as AP802, AP804, AP806, AP810 and AP814. The carbon dioxide and hydrogen sulfide are absorbed by the amine while the hydrogen passes through first amine scrubber zone and into line to be recycled to reaction zone 304. The amine is regenerated and the carbon dioxide and hydrogen sulfide are released and removed in line 362. Within the first amine absorber zone, regenerated amine may be recycled for use again. The released carbon dioxide and hydrogen sulfide in line 362 are passed through second amine scrubber zone 358 which contains an amine selective to hydrogen sulfide, but not selective to carbon dioxide. Again, suitable amines are available from DOW and from BASF, and in one embodiment the amines are a promoted or activated MDEA. Suitable amines for the second amine absorber zone from DOW include the UCARSOL™ HS series solvents such as HS101, HS102, HS103, HS104, HS115. Therefore the carbon dioxide passes through second amine scrubber zone 358 and into line 366. The amine may be regenerated which releases the hydrogen sulfide into line 360. At least a portion of the hydrogen sulfide in line 360 may be recycled to the reaction zone 304, possibly in measured controlled amount. Excess hydrogen sulfide may be directed to a Claus plant. Regenerated amine is reused. Conditions for the first amine solution absorber zone includes from about 30 to about 60° C. and a pressure in the range of about 3447 kPa (500 psia) to about 6895 kPa (1000 psia) and conditions for the second amine solution absorber zone (258) includes from about 20 to about 35° C. and a pressure in the range of about 138 kPa (20 psia) to about 241 kPa (35 psia).
In another embodiment, the amine solution absorber zone 356 may contain the amine solution of U.S. Pat. No. 4,710,362 which selectively separates only the carbon dioxide and allows the hydrogen sulfide to pass with the hydrogen into recycle line 376. In this embodiment, the second amine absorber zone 358 is not necessary.
It should be noted in
Turning to
Deoxygenation reactor effluent stream 406 after one or more optional heat exchanges, is directed to hot high pressure hydrogen stripper 408. Hydrogen-rich mixture stream 22 is also introduced to hot high pressure hydrogen stripper 408. In hot high pressure hydrogen stripper 408, the gaseous component of deoxygenation reactor effluent 406 is stripped from the liquid component of deoxygenation reactor effluent 406 using hydrogen-rich stream 22. The gaseous component comprising hydrogen, vaporous water, carbon monoxide, carbon dioxide and possibly some hydrogen sulfide, is separated into hot high pressure hydrogen stripper overhead stream 414. The remaining liquid component of deoxygenation reactor effluent 406 comprising primarily normal paraffins having a carbon number from about 8 to about 24 with a cetane number of about 60 to about 100 is removed as hot high pressure hydrogen stripper bottoms 412.
A portion of hot high pressure hydrogen stripper bottoms forms recycle stream 413 and is combined with renewable feedstock combined stream 420. Another portion of recycle stream 413 may be routed directly to deoxygenation reactor 404 and introduced at interstage locations such as between beds 404a and 404b and/or between beds 404b and 404c in order, for example, to aid in temperature control and/or to enhance the soluble hydrogen levels in the reactants. Hot steam could be injected through line 415 to combine with stream 313 to provide reactant water for the water gas shift reaction. When the optional water gas shift reactor 20 is bypassed or not included, stream 22 contains large amount of carbon monoxide, which is a regulated air pollutant. In addition, carbon monoxide is considered a fuel, and releasing carbon monoxide to the atmosphere is a waste of energy. The reactor 404 can assume the function of carrying out the water gas shift reaction, in addition to the hydrogenation and de-oxygenation reactions. Line 415 can provide the reactant water needed for water gas shift reaction. The water vapor can also help suppress the hydrodeoxygenation reaction, which tends to consume more hydrogen.
The remainder of hot high pressure hydrogen stripper bottoms in stream 412 is routed to isomerization zone 416 where it contacts an isomerization catalyst to convert normal paraffins to branched paraffins. Stream 412 may be heat exchanged with isomerization reactor effluent 422. The product of the isomerization reactor containing a gaseous portion of hydrogen and a branched-paraffin-rich liquid portion is removed in line 422, and after optional heat exchange with stream 412, is introduced into hydrogen separator 426. The overhead stream 428 from hydrogen separator 426 contains primarily hydrogen which may be recycled back to hot high pressure hydrogen stripper 408. Bottom stream 429 from hydrogen separator 426 is air cooled using air cooler 432 and introduced into product separator 434. In product separator 434 the gaseous portion of the stream comprising hydrogen, carbon monoxide, hydrogen sulfide and carbon dioxide phase separate and are removed in stream 436 while the liquid hydrocarbon portion of the stream is removed in stream 438. A liquid water byproduct stream 440 may also be removed from product separator 434. Stream 438 is introduced to product recovery column 442 where components having higher relative volatilities are separated into stream 444 with the remainder, the diesel range components, being withdrawn from product recovery column 442 in line 446.
The vapor stream 436 from product separator 434 contains the gaseous portion of the isomerization effluent which comprises at least hydrogen, carbon monoxide, hydrogen sulfide and carbon dioxide and is directed to a system of at least one amine solution absorber to separate carbon dioxide and hydrogen sulfide from the vapor stream. Because of the cost of hydrogen, it is desirable to recycle the hydrogen to deoxygenation reactor 404, but it is not desirable to circulate the carbon dioxide or too much of an excess of sulfur containing components. In order to separate sulfur containing components and carbon dioxide from the hydrogen, vapor stream 436 is passed through a system of at least two amine solution absorbers, also called scrubbers, starting with the first amine solution absorber zone 456. The carbon dioxide and hydrogen sulfide are absorbed by the amine while the hydrogen passes through first amine absorber zone and into line 476 to be recycled to reaction zone 404. The amine is regenerated and the carbon dioxide and hydrogen sulfide are released and removed in line 462. Within the first amine solution absorber zone, regenerated amine may be recycled for use again. The released carbon dioxide and hydrogen sulfide in line 462 are passed through second amine solution absorber zone 458 which contains an amine selective to hydrogen sulfide, but not selective to carbon dioxide. Therefore the carbon dioxide passes through second amine absorber zone 458 and into line 466. The amine may be regenerated which releases the hydrogen sulfide into line 460. At least a portion of the hydrogen sulfide in line 460 may be recycled (not shown) to the reaction zone 404, possibly in measured controlled amount. Excess hydrogen sulfide may be directed to a Claus plant. Regenerated amine is reused.
Conditions for the first amine solution absorber zone include from about 20 to about 60° C. and a pressure in the range of about 3447 kPa (500 psia) to about 6895 kPa (1000 psia). Conditions for the second amine solution absorber zone includes from about 20 to about 60° C. and a pressure in the range of about 138 kPa (20 psia) to about 241 kPa (35 psia).
In another embodiment, the amine solution absorber zone 456 may contain the amine solution of U.S. Pat. No. 4,710,362 which selectively separates only the carbon dioxide and allows the hydrogen sulfide to pass with the hydrogen into recycle line 476. In this embodiment, the second amine absorber zone 458 is not necessary.
