Patent Application
20120289752
The present application claims the benefit of European Patent Application No. 11163542.1, filed Apr. 21, 2011 the entire disclosure of which is hereby incorporated by reference.
The invention relates to a process for converting a solid biomass material, products thereof and a process for producing a biofuel and/or biochemical and products thereof.
With the diminishing supply of crude mineral oil, use of renewable energy sources is becoming increasingly important for the production of liquid fuels. These fuels from renewable energy sources are often referred to as biofuels.
Biofuels derived from non-edible renewable energy sources, such as cellulosic materials, are preferred as these do not compete with food production. These biofuels are also referred to as second generation biofuels, renewable or advanced biofuels. Most non-edible renewable energy sources, however, are solid biomass materials that are cumbersome to convert into liquid fuels.
For example, the process described in WO 2010/062611 for converting solid biomass to hydrocarbons requires three catalytic conversion steps. First the solid biomass is contacted with a catalyst in a first riser operated at a temperature in the range of from about 50 to about 200° C. to produce a first biomass-catalyst mixture and a first product comprising hydrocarbons (referred to as pretreatment). Hereafter the first biomass-catalyst mixture is charged to a second riser operated at a temperature in the range of from about 200° to about 400° C. to thereby produce a second biomass-catalyst mixture and a second product comprising hydrocarbons (referred to as deoxygenating and cracking); and finally the second biomass-catalyst mixture is charged to a third riser operated at a temperature greater than about 450° C. to thereby produce a spent catalyst and a third product comprising hydrocarbons. The last step is referred to as conversion to produce the fuel or specialty chemical product. WO 2010/062611 mentions the possibility of preparing the biomass for co-processing in conventional petroleum refinery units. The process of WO 2010/062611, however, is cumbersome in that three steps are needed, each step requiring its own specific catalyst.
WO2010/135734 describes a method for co-processing a biomass feedstock and a refinery feedstock in a refinery unit comprising catalytically cracking the biomass feedstock and the refinery feedstock in a refinery unit comprising a fluidized reactor, wherein hydrogen is transferred from the refinery feedstock to carbon and oxygen of the biomass feedstock. In one of the embodiments WO2010/135734 the biomass feedstock comprises a plurality of solid biomass particles having an average size between 50 and 1000 microns. In passing, it is further mentioned that solid biomass particles can be pre-processed to increase brittleness, susceptibility to catalytic conversion (e.g. by roasting, toasting, and/or torrefication) and/or susceptibility to mixing with a petrochemical feedstock.
It has now for the first time been recognized that a drawback of the processes in the prior art and in specific the process as described in WO2010/135734 is that the product(s) obtained from the fluidized reactor for catalytic cracking will contain a much higher concentration of oxygen-containing-hydrocarbons than the product obtained by catalytic cracking of a conventional feed. These oxygen-containing hydrocarbons, such as for example ethers, esters, acids and/or alcohols may have several disadvantages when present in a biofuel and/or biochemical. For example, oxygen containing compounds such as ethers may under certain circumstances lead to peroxide formation when contacted with air, which peroxide formation may increase explosion risks downstream of the reactor. Acids present in the product may lead to corrosion downstream of the reactor or, if such acids are included in a biofuel, corrosion in the engines of the cars of the end-users. Alcohols like phenols present in the product may be toxic for downstream waste water cleaning units. In addition, also any ethers or esters present may be undesirable for the downstream waste water cleaning units as they may lead to groundwater contamination in case of a spill.
It would be an advancement in the art to provide a process for converting a solid biomass material into a biofuel and/or a biochemical component, not having the above drawbacks or disadvantage(s).
Such a process has been achieved with the process according to the invention.
Accordingly one embodiment provides a process for converting a solid biomass material comprising
a) contacting the solid biomass material and a fluid hydrocarbon feed with a catalytic cracking catalyst at a temperature of more than 400° C. in a catalytic cracking reactor to produce at least one cracked product;
b) fractionating the cracked product produced in step a) to produce at least one product fraction;
c) hydrodeoxygenating at least one product fraction produced in step b) to produce at least one hydrodeoxygenated product.
Some of these hydrodeoxygenated products are novel and inventive in itself. The present invention therefore also provides a hydrodeoxygenated product composition comprising
from equal to or more than 0 wt % to equal to or less than 60 wt % olefins;
from equal to or more than 0 wt % to equal to or less than 1 wt % oxygen-containing hydrocarbons;
from equal to or more than 5 wt % to equal to or less than 80 wt % paraffins;
from equal to or more than 1 wt % to equal to or less than 60 wt % cycloparaffins;
from equal to or more than 1 wt % to equal to or less than 60 wt % aromatics, based on the total weight of the composition;
which composition comprises in the range from equal to or more than 0.02 wt % to equal to or less than 50 wt % of bio-carbon, based on the total weight of carbon present in the composition.
The hydrodeoxygenated product can conveniently be used as a biofuel component and/or a biochemical component or can be converted into a biofuel component and/or a biochemical component.
The invention further also provides a process for the preparation of a biofuel respectively a biochemical comprising a biofuel component respectively a biochemical component wherein the biofuel component respectively the biochemical component comprises one or more hydrodeoxygenated products obtained in a process as described above or wherein the biofuel component respectively the biochemical component is derived from one or more hydrodeoxygenated products obtained in a process as described above.
The present invention also provides a biofuel composition comprising
i) a conventional fuel component
ii) a biofuel component comprising
hydrodeoxygenated product composition comprising
from equal to or more than 0 wt % to equal to or less than 60 wt % olefins;
from equal to or more than 0 wt % to equal to or less than 1 wt % oxygen-containing hydrocarbons;
from equal to or more than 5 wt % to equal to or less than 80 wt % paraffins;
from equal to or more than 1 wt % to equal to or less than 60 wt % cycloparaffins;
from equal to or more than 1 wt % to equal to or less than 60 wt % aromatics, based on the total weight of the composition;
wherein the biofuel component comprises in the range from equal to or more than 0.02 wt % to equal to or less than 50 wt % of bio-carbon, based on the total weight of carbon present in the composition.
The process according to the invention allows one to prepare a biofuel and/or biochemical component via catalytic cracking of a solid biomass material, wherein the biofuel and/or biochemical component has a minimal concentration of oxygen-containing hydrocarbons.
The fluid hydrocarbon co-feed provides hydrogen, which hydrogen can advantageously be used in the removal of oxygen by hydrogen transfer during the catalytic cracking reaction and/or in the hydrodeoxygenation of one or more of the product fraction(s) of step b). Without wishing to be bound by any kind of theory it is believed that the fluid hydrocarbon co-feed therefore assists in minimization of the formation of oxygen-containing hydrocarbons.
The process according to the invention can be simple and may require a minimum of processing steps to convert a solid biomass material to a biofuel component or biochemical component that is low in oxygen content. Such biofuel component may be fully fungible.
Furthermore the process according to the invention may be easily implemented in existing refineries.
In addition, the process according to the invention may not need any complicated actions, for example it may not need a pre-mixed composition of the solid biomass material and the catalyst.
The process according to the invention therefore also provides a more direct route via catalytic cracking of solid biomass material to second generation, renewable or advanced, biofuels and/or biochemicals.
In step a) of the process a solid biomass material and a fluid hydrocarbon feed are contacted with a catalytic cracking catalyst at a temperature of more than 400° C. in a catalytic cracking reactor to produce one or more cracked products.
By a solid biomass material is herein understood a solid material obtained from a renewable source. By a renewable source is herein understood a composition of matter of biological origin as opposed to a composition of matter obtained or derived from petroleum, natural gas or coal. Without wishing to be bound by any kind of theory it is believed that such material obtained from a renewable source may contain carbon-14 isotope in an abundance of about 0.0000000001%, based on total moles of carbon.
Preferably the renewable source is a composition of matter of cellulosic or lignocellulosic origin. Any solid biomass material may be used in the process of the invention. In a preferred embodiment the solid biomass material is not a material used for food production. Examples of preferred solid biomass materials include aquatic plants and algae, agricultural waste and/or forestry waste and/or paper waste and/or plant material obtained from domestic waste.
Preferably the solid biomass material contains cellulose and/or lignocellulose. Examples of suitable cellulose- and/or lignocellulose-containing materials include agricultural wastes such as corn stover, soybean stover, corn cobs, rice straw, rice hulls, oat hulls, corn fibre, cereal straws such as wheat, barley, rye and oat straw; grasses; forestry products and/or forestry residues such as wood and wood-related materials such as sawdust; waste paper; sugar processing residues such as bagasse and beet pulp; or mixtures thereof. More preferably the solid biomass material is selected from the group consisting of wood, sawdust, straw, grass, bagasse, corn stover and/or mixtures thereof.
The solid biomass material may have undergone drying, torrefaction, steam explosion, particle size reduction, densification and/or pelletization before being contacted with the catalyst, to allow for improved process operability and economics.
Preferably the solid biomass material in step a) is a torrefied solid biomass material. In a preferred embodiment the process according to the invention comprises a step of torrefying the solid biomass material at a temperature of more than 200° C. to obtain a torrefied solid biomass material that can be contacted with the catalytic cracking catalyst in step a). The words torrefying and torrefaction are used interchangeable herein.
By torrefying or torrefaction is herein understood the treatment of the solid biomass material at a temperature in the range from equal to or more than 200° C. to equal to or less than 350° C. in the essential absence of a catalyst and in an oxygen-poor, preferably an oxygen-free, atmosphere. By an oxygen-poor atmosphere is understood an atmosphere containing equal to or less than 15 vol. % oxygen, preferably equal to or less than 10 vol % oxygen and more preferably equal to or less than 5 vol % oxygen. By an oxygen-free atmosphere is understood that the torrefaction is carried out in the essential absence of oxygen.
Torrefying of the solid biomass material is preferably carried out at a temperature of more than 200° C., more preferably at a temperature equal to or more than 210° C., still more preferably at a temperature equal to or more than 220° C., yet more preferably at a temperature equal to or more than 230° C. In addition torrefying of the solid biomass material is preferably carried out at a temperature less than 350° C., more preferably at a temperature equal to or less than 330° C., still more preferably at a temperature equal to or less than 310° C., yet more preferably at a temperature equal to or less than 300° C.
Torrefaction of the solid biomass material is preferably carried out in the essential absence of oxygen. More preferably the torrefaction is carried under an inert atmosphere, containing for example inert gases such as nitrogen, carbon dioxide and/or steam; and/or under a reducing atmosphere in the presence of a reducing gas such as hydrogen, gaseous hydrocarbons such as methane and ethane or carbon monoxide.
The torrefying step may be carried out at a wide range of pressures. Preferably, however, the torrefying step is carried out at atmospheric pressure (about 1 bar, corresponding to about 0.1 MegaPascal).
The torrefying step may be carried out batchwise or continuously. If operated batchwise, a torrefaction reactor may be filled with solid biomass material, whereafter the solid biomass material in the torrefaction reactor may be heated at the torrefaction temperature for a time period in the range from 1 minute to 12 hours, more preferably for a period in the range from 30 minutes to 8 hours and most preferably for a period in the range from 1 to 6 hours. Hereafter the torrefaction reactor may be cooled down and emptied to start a new cycle.
If operated continuously, for example the TORSPYD (trademark) process of Thermya may be used, wherein a flow of solid biomass material flows from top to bottom in a reactor column, counter-currently to a flow of gas flowing from bottom to top of the reactor column. The temperature of the reactor column gradually increases from top to bottom. Residence time for the solid biomass material in the torrefaction reactor may lie in the range from equal to or more than 0.5 minute, more preferably equal to or more than 5 minutes and most preferably equal to or more than 15 minutes to equal to or less than 2 hours, more preferably equal to or less than 1 hour and most preferably equal to or less than 45 minutes.
The torrefied solid biomass material has a higher energy density, a higher mass density and greater flowability, making it easier to transport, pelletize and/or store. Being more brittle, it can be easier reduced into smaller particles.
Preferably the torrefied solid biomass material has an oxygen content in the range from equal to or more than 10 wt %, more preferably equal to or more than 20 wt % and most preferably equal to or more than 30 wt % oxygen, to equal to or less than 60 wt %, more preferably equal to or less than 50 wt %, based on total weight of dry matter.
During torrefaction of the solid biomass material torrefaction gases can be produced. These torrefaction gases can contain carbon monoxide and carbon dioxide but also volatile fuels such as for example methane, ethane, ethene and/or methanol. In a preferred embodiment according to this invention, these volatile fuels are retrieved from the torrefaction gases and recycled to the process as a fuel to provide at least part of the heat for the torrefaction and/or the cracking in step (a). In a further embodiment carbon monoxide and/or carbon dioxide may be retrieved from the torrefaction gases and recycled to provide the inert or reducing atmosphere for torrefaction.
In a further preferred embodiment, any torrefying or torrefaction step further comprises drying the solid biomass material before such solid biomass material is torrefied. In such a drying step, the solid biomass material is preferably dried until the solid biomass material has a moisture content in the range of equal to or more than 0.1 wt % to equal to or less than 25 wt %, more preferably in the range of equal to or more than 5 wt % to equal to or less than 20 wt %, and most preferably in the range of equal to or more than 5 wt % to equal to or less than 15 wt %. For practical purposes moisture content can be determined via ASTM E1756-01 Standard Test method for Determination of Total solids in Biomass. In this method the loss of weight during drying is a measure for the original moisture content.
Preferably the solid biomass material in step a) is a micronized solid biomass material. By a micronized solid biomass material is herein understood a solid biomass material that has a particle size distribution with a mean particle size in the range from equal to or more than 5 micrometer to equal to or less than 5000 micrometer, as measured with a laser scattering particle size distribution analyzer. In a preferred embodiment the process according to the invention comprises a step of reducing the particle size of the solid biomass material, optionally before or after such solid biomass material is torrefied. Such a particle size reduction step may for example be especially advantageous when the solid biomass material comprises wood or torrefied wood. The particle size of the, optionally torrefied, solid biomass material can be reduced in any manner known to the skilled person to be suitable for this purpose. Suitable methods for particle size reduction include crushing, grinding and/or milling. The particle size reduction may for example be achieved by means of a ball mill, hammer mill, (knife) shredder, chipper, knife grid, or cutter.
Preferably the solid biomass material has a particle size distribution where the mean particle size lies in the range from equal to or more than 5 micrometer (micron), more preferably equal to or more than 10 micrometer, even more preferably equal to or more than 20 micrometer, and most preferably equal to or more than 100 micrometer to equal to or less than 5000 micrometer, more preferably equal to or less than 1000 micrometer and most preferably equal to or less than 500 micrometer.
In one especially preferred embodiment the solid biomass material has a particle size distribution where the mean particle size is equal to or more than 100 micrometer to avoid blocking of pipelines and/or nozzles. Most preferably the solid biomass material has a particle size distribution where the mean particle size is equal to or less than 3000 micrometer to allow easy injection into the riser reactor.
In an alternative preferred embodiment the solid biomass material has a particle size distribution where the mean particle size is equal to or more than 2000 micrometer, more preferably equal to or more than 2500 micrometer, most preferably equal to or more than 3000 micrometer. As explained in more detail herein below, the process according to the invention can be carried out such that a longer residence time in the catalytic cracking reactor is obtained for the solid biomass material. Such longer residence times in turn allow one to advantageously use solid biomass material with larger particles. The solid biomass material with larger particles will require less energy to prepare. For practical purposes it may be preferred that the solid biomass material in this case has a particle size distribution where the mean particle size is equal to or less than 2 cm, more preferably equal to or less than 1 cm, most preferably equal to or less than 5000 micrometer. The process of this embodiment may be novel and inventive of its own and therefore the present invention further provides a process for converting a solid biomass material comprising contacting the solid biomass material and a fluid hydrocarbon feed with a catalytic cracking catalyst at a temperature of more than 400° C. in a catalytic cracking reactor to produce one or more cracked products, wherein the solid biomass material has a particle size distribution with a mean particle size of equal to or more than 2000 micrometer. Preferences for such a process are further as described herein above and herein below.
For practical purposes the particle size distribution and mean particle size of the solid biomass material can be determined with a Laser Scattering Particle Size Distribution Analyzer, preferably a Horiba LA950, according to the ISO 13320 method titled “Particle size analysis—Laser diffraction methods”.
Hence, preferably the process of the invention comprises a step of reducing the particle size of the solid biomass material, optionally before and/or after torrefaction, to generate a particle size distribution having a mean particle size in the range from equal to or more than 5, more preferably equal to or more than 10 micron, and most preferably equal to or more than 20 micron, to equal to or less than 2 cm, more preferably to equal to or less than 5000 micrometer (micron), more preferably equal to or less than 1000 micrometer and most preferably equal to or less than 500 micrometer to produce a micronized, optionally torrefied, solid biomass material.
In an optional embodiment the particle size reduction of the, optionally torrefied, solid biomass material is carried out whilst having the solid biomass material suspended in a hydrocarbon-containing liquid, to improve processibility and/or avoid dusting. In one preferred embodiment such a suspension of solid biomass particles in a hydrocarbon-containing liquid is prepared by a process comprising a first particle size reduction step wherein the particle size of a solid biomass material is reduced to produce a first particulate product comprising solid biomass particles; a mixing step, wherein the first particulate product is suspended in a hydrocarbon-containing liquid to produce a suspended first particulate product comprising solid biomass particles suspended in the hydrocarbon-containing liquid; and a second particle size reduction step, wherein the particle size of the suspended first particulate product is further reduced to produce a suspended second particulate product comprising solid biomass particles suspended in the hydrocarbon-containing liquid.
Preferably at least 80 wt % of the first particulate product has a particle size of equal to or less than 300 micrometer and at least 80 wt % of the second particulate product has a particle size of equal to or less than 100 micrometer.
Preferably the hydrocarbon-containing liquid comprises straight run (atmospheric) gas oils, flashed distillate, vacuum gas oils (VGO), coker gas oils, gasoline, naphtha, diesel, kerosene, atmospheric residue (“long residue”) and vacuum residue (“short residue”) and/or mixtures thereof. Most preferably the hydrocarbon-containing liquid comprises gasoline, naphtha, diesel, kerosene, and/or mixtures thereof.
In one embodiment the fluid hydrocarbon co-feed as described herein below is used as hydrocarbon-containing liquid.
In a preferred embodiment the, optionally micronized and optionally torrefied, solid biomass material is dried before being supplied to the riser reactor. Hence, if the solid biomass material is torrefied, it may be dried before and/or after torrefaction. If dried before use as a feed to the riser reactor, the solid biomass material is preferably dried at a temperature in the range from equal to or more than 50° C. to equal to or less than 200° C., more preferably in the range from equal to or more than 80° C. to equal to or less than 150° C. The, optionally micronized and/or torrefied, solid biomass material is preferably dried for a period in the range from equal to or more than 30 minutes to equal to or less than 2 days, more preferably for a period in the range from equal to or more than 2 hours to equal to or less than 24 hours.
In addition to the, optionally micronized and/or optionally torrefied, solid biomass material also a fluid hydrocarbon feed (herein also referred to as fluid hydrocarbon co-feed) is contacted with the catalytic cracking catalyst in the catalytic cracking reactor.
By a hydrocarbon feed is herein understood a feed that contains one or more hydrocarbon compounds. By hydrocarbon compounds are herein understood compounds that contain or preferably consist of both hydrogen and carbon. By a fluid hydrocarbon feed is herein understood a hydrocarbon feed that is not in a solid state. The fluid hydrocarbon co-feed is preferably a liquid hydrocarbon co-feed, a gaseous hydrocarbon co-feed, or a mixture thereof. The fluid hydrocarbon co-feed can be fed to a catalytic cracking reactor (preferably a riser reactor) in an essentially liquid state, in an essentially gaseous state or in a partially liquid-partially gaseous state. When entering the catalytic cracking reactor in an essentially or partially liquid state, the fluid hydrocarbon co-feed preferably vaporizes upon entry and preferably is contacted in the gaseous state with the catalytic cracking catalyst and/or the solid biomass material.
The fluid hydrocarbon feed can be any non-solid hydrocarbon feed known to the skilled person to be suitable as a feed for a catalytic cracking reactor. The fluid hydrocarbon feed can for example be obtained from a conventional crude oil (also sometimes referred to as a petroleum oil or mineral oil), an unconventional crude oil (that is, oil produced or extracted using techniques other than the traditional oil well method) or a renewable oil (that is, oil derived from a renewable source, such as pyrolysis oil, vegetable oil or the products of a biomass liquefaction process), a Fisher Tropsch oil (sometimes also referred to as a synthetic oil) and/or a mixture of any of these.
In one embodiment the fluid hydrocarbon feed is derived from a, preferably conventional, crude oil. Examples of conventional crude oils include West Texas Intermediate crude oil, Brent crude oil, Dubai-Oman crude oil, Arabian Light crude oil, Midway Sunset crude oil or Tapis crude oil.
More preferably the fluid hydrocarbon feed comprises a fraction of a, preferably conventional, crude oil or renewable oil. Preferred fluid hydrocarbon feeds include straight run (atmospheric) gas oils, flashed distillate, vacuum gas oils (VGO), coker gas oils, diesel, gasoline, kerosene, naphtha, liquefied petroleum gases, atmospheric residue (“long residue”) and vacuum residue (“short residue”) and/or mixtures thereof. Most preferably the fluid hydrocarbon feed comprises a long residue, a vacuum gas oil or a mixture thereof.
In one embodiment the fluid hydrocarbon feed preferably has a 5 wt % boiling point, as measured by means of distillation as based on ASTM D86 titled “Standard Test Method for Distillation of Petroleum Products at Atmospheric Pressure”, respectively as measured by ASTM D1160 titled “Standard Test Method for Distillation of Petroleum Products at Reduced Pressure”, at a pressure of 1 bar absolute (0.1 MegaPascal), of equal to or more than 100° C., more preferably equal to or more than 150° C. An example of such a fluid hydrocarbon feed is vacuum gas oil.
In a second embodiment the fluid hydrocarbon feed preferably has a 5 wt % boiling point, as measured by means of distillation based on ASTM D86 titled “Standard Test Method for Distillation of Petroleum Products at Atmospheric Pressure”, respectively as measured by ASTM D1160 titled “Standard Test Method for Distillation of Petroleum Products at Reduced Pressure”, at a pressure of 1 bar absolute (0.1 MegaPascal), of equal to or more than 200° C., more preferably equal to or more than 220° C., most preferably equal to or more than 240° C. An example of such a fluid hydrocarbon feed is long residue.
In a further preferred embodiment equal to or more than 70 wt %, preferably equal to or more than 80 wt %, more preferably equal to or more than 90 wt % and still more preferably equal to or more than 95 wt % of the fluid hydrocarbon feed boils in the range from equal to or more than 150° C. to equal to or less than 600° C., as measured by means of a distillation by ASTM D86 titled “Standard Test Method for Distillation of Petroleum Products at Atmospheric Pressure”, respectively as measured by ASTM D1160 titled “Standard Test Method for Distillation of Petroleum Products at Reduced Pressure”, at a pressure of 1 bar absolute (0.1 MegaPascal).
The composition of the fluid hydrocarbon feed may vary widely. The fluid hydrocarbon feed may for example contain paraffins, naphthenes, olefins and/or aromatics.
Preferably the fluid hydrocarbon feed comprises in the range from equal to or more than 50 wt %, more preferably from equal to or more than 75 wt %, and most preferably from equal to or more than 90 wt % to equal to or less than 100 wt % of compounds consisting only of carbon and hydrogen, based on the total weight of the fluid hydrocarbon feed.
Preferably the fluid hydrocarbon feed comprises equal to or more than 1 wt % paraffins, more preferably equal to or more than 5 wt % paraffins, and most preferably equal to or more than 10 wt % paraffins, and preferably equal to or less than 100 wt % paraffins, more preferably equal to or less than 90 wt % paraffins, and most preferably equal to or less than 30 wt % paraffins, based on the total fluid hydrocarbon feed. By paraffins both normal-, cyclo- and branched-paraffins are understood.
In another embodiment the fluid hydrocarbon feed comprises or consists of a paraffinic fluid hydrocarbon feed. By a paraffinic fluid hydrocarbon feed is herein understood a fluid hydrocarbon feed comprising in the range from at least 50 wt % of paraffins, preferably at least 70 wt % of paraffins, and most preferably at least 90 wt % paraffins, up to and including 100 wt % paraffins, based on the total weight of the fluid hydrocarbon feed.
For practical purposes the paraffin content of all fluid hydrocarbon feeds having an initial boiling point of at least 260° C. can be measured by means of ASTM method D2007-03 titled “Standard test method for characteristic groups in rubber extender and processing oils and other petroleum-derived oils by clay-gel absorption chromatographic method”, wherein the amount of saturates will be representative for the paraffin content. For all other fluid hydrocarbon feeds the paraffin content of the fluid hydrocarbon feed can be measured by means of comprehensive multi-dimensional gas chromatography (GC×GC), as described in P. J. Schoenmakers, J. L. M. M. Oomen, J. Blomberg, W. Genuit, G. van Velzen, J. Chromatogr. A, 892 (2000) p. 29 and further.
Examples of such paraffinic fluid hydrocarbon co-feeds include so-called Fischer-Tropsch derived hydrocarbon streams such as described in WO2007/090884 and herein incorporated by reference, or a hydrogen rich feed like hydrotreater product or hydrowax. By Hydrowax is understood the bottoms fraction of a hydrocracker. Examples of hydrocracking processes which may yield a bottoms fraction that can be used as fluid hydrocarbon co-feed, are described in EP-A-699225, EP-A-649896, WO-A-97/18278, EP-A-705321, EP-A-994173 and U.S. Pat. No. 4,851,109 and herein incorporated by reference.
By “Fischer-Tropsch derived hydrocarbon stream” is meant that the hydrocarbon stream is a product from a Fischer-Tropsch hydrocarbon synthesis process or derived from such product by a hydroprocessing step, i.e. hydrocracking, hydro-isomerisation and/or hydrogenation.
The Fischer-Tropsch reaction converts carbon monoxide and hydrogen into longer chain, usually paraffinic, hydrocarbons:
n(CO+2H2)=(—CH2—)n+nH2O+heat,
in the presence of an appropriate catalyst and preferably at elevated temperature, for example 125 to 300° C., preferably 175 to 250° C., and elevated pressure, for example 5 to 100 bar (0.5 to 10 MegaPascal), preferably 12 to 80 bar (1.2 to 8.0 MegaPascal).
The carbon monoxide and hydrogen is typically derived from a hydrocarbonaceous feedstock by partial oxidation. Suitable hydrocarbonaceous feedstocks for such partial oxidation include gaseous hydrocarbons such as natural gas or methane, coal, biomass, or residual fractions from crude oil distillation.
The Fischer-Tropsch derived hydrocarbon stream may suitably be a so-called syncrude as described in for example GB-A-2386607, GB-A-2371807 or EP-A-0321305. Other suitable Fischer-Tropsch hydrocarbon streams may be hydrocarbon fractions boiling in the naphtha, kerosene, gas oil, or wax range, as obtained from the Fischer-Tropsch hydrocarbon synthesis process, optionally followed by a hydroprocessing step.
Preferably, the Fischer-Tropsch hydrocarbon stream product has been obtained by hydroisomerisation of hydrocarbons directly obtained in the Fischer-Tropsch hydrocarbon synthesis reaction. The use of a hydro-isomerised hydrocarbon fraction is advantageous because it contributes to a high yield in gasoline due to the high content of iso-paraffins in said fraction. A hydro-isomerised fraction boiling in the kerosene or gas oil range may suitable be used as the Fischer-Tropsch derived hydrocarbon stream. Preferably, however, a higher boiling hydro-isomerised fraction is used.
A particularly suitable hydro-isomerised hydrocarbon fraction is a fraction which has a T10 wt % boiling point of between 350 and 450° C. and a T90 wt % of between 450 and 600° C. and a wax content of between 5 and 60 wt %. Such fraction is typically referred to as waxy raffinate. Preferably, the wax content is between 5 and 30 wt %. The wax content is measured by solvent dewaxing at −27° C. in a 50/50 vol./vol. mixture of methyl ethyl ketone and toluene. Examples of such a hydrocarbon streams are the commercially available Waxy Raffinate product as is marketed by Shell MDS (Malaysia) and the waxy raffinate product as obtained by the process described in WO-A-02/070630 or in EP-B-0668342.
In an especially preferred embodiment the fluid hydrocarbon co-feed contains:
a fraction of a crude oil, such as for example (atmospheric) gas oils, flashed distillate, vacuum gas oils (VGO), coker gas oils, atmospheric residue (“long residue”) and vacuum residue (“short residue”); in combination with
a paraffinic fluid hydrocarbon co-feed as described above.
The weight ratio of the solid biomass material to fluid hydrocarbon co-feed may vary widely. For ease of co-processing the weight ratio of fluid hydrocarbon co-feed to solid biomass material is preferably equal to or more than 50 to 50 (5:5), more preferably equal to or more than 70 to 30 (7:3), still more preferably equal to or more than 80 to 20 (8:2), even still more preferably equal to or more than 90 to 10 (9:1). For practical purposes the weight ratio of fluid hydrocarbon co-feed to solid biomass material is preferably equal to or less than 99.9 to 0.1 (99.9:0.1), more preferably equal to or less than 95 to 5 (95:5). The fluid hydrocarbon co-feed and the solid biomass material are preferably being fed to a catalytic cracking reactor in a weight ratio within the above ranges.
The amount of solid biomass material, based on the total weight of solid biomass material and fluid hydrocarbon co-feed present in a feed to a catalytic cracking reactor, is preferably equal to or less than 30 wt %, more preferably equal to or less than 20 wt %, most preferably equal to or less than 10 wt % and even more preferably equal to or less than 5 wt %. For practical purposes the amount of solid biomass material present, based on the total weight of solid biomass material and fluid hydrocarbon co-feed present in a feed to a catalytic cracking reactor, is preferably equal to or more than 0.1 wt %, more preferably equal to or more than 1 wt %.
In an especially preferred process the total feed to step (a) comprises:
1) from equal to or more than 0 wt % to equal to or less than 99 wt %, preferably from equal to or more than 0 wt % to equal to or less than 20 wt % of a paraffinic fluid hydrocarbon co-feed;
2) from equal to or more than 0 wt % to equal to or less than 99 wt %, preferably from equal to or more than 60 wt % to equal to or less than 80 wt % of a fraction of a crude oil, such as for example (atmospheric) gas oils, flashed distillate, naphtha, diesel, kerosene, liquefied petroleum gas, vacuum gas oils (VGO), coker gas oils, atmospheric residue (“long residue”) and vacuum residue (“short residue”); and
3) from equal to or more than 1 wt % to equal to or less than 35 wt %, preferably from equal to or more than 1 wt % to equal to or less than 20 wt % of a solid biomass material or a part thereof as described herein.
In a preferred embodiment the fluid hydrocarbon co-feed comprises equal to or more than 8 wt % elemental hydrogen, more preferably more than 12 wt % elemental hydrogen, based on the total fluid hydrocarbon co-feed on a dry basis (i.e. on a water-free basis). A high content of elemental hydrogen, such as a content of equal to or more than 8 wt %, allows the hydrocarbon co-feed to act as a cheap hydrogen donor in the catalytic cracking process. A particularly preferred fluid hydrocarbon co-feed having an elemental hydrogen content of equal to or more than 8 wt % is Fischer-Tropsch derived waxy raffinate. Such Fischer-Tropsch derived waxy raffinate may for example comprise about 85 wt % of elemental carbon and 15 wt % of elemental hydrogen.
Without wishing to be bound by any kind of theory it is further believed that a higher weight ratio of fluid hydrocarbon co-feed to solid biomass material enables more upgrading of the solid biomass material by hydrogen transfer reactions.
Preferably step (a) is carried out in a catalytic cracking unit, more preferably in a fluidized catalytic cracking (FCC) unit. Preferably the catalytic cracking unit comprises at least a catalytic cracking reactor and a catalyst regenerator.
The catalytic cracking reactor used in step (a) can be any catalytic cracking reactor known in the art to be suitable for the purpose, including for example a fluidized bed reactor or a riser reactor. Most preferably the catalytic cracking reactor is a riser reactor.
The fluid hydrocarbon co-feed and the, optionally micronized and/or optionally torrefied, solid biomass material can be mixed prior to entry into a catalytic cracking reactor or they can be added separately, at the same location or at different locations to the catalytic cracking reactor.
In one embodiment the fluid hydrocarbon co-feed and the, optionally micronized and/or torrefied, solid biomass material are not mixed together prior to entry into a catalytic cracking reactor. In this embodiment the fluid hydrocarbon co-feed and the solid biomass material may be fed simultaneously (that is at one location) to the catalytic cracking reactor, and optionally mixed upon entry of the catalytic cracking reactor; or, alternatively, the fluid hydrocarbon co-feed and the solid biomass material may be added separately (at different locations) to the catalytic cracking reactor. Catalytic cracking reactors, and especially riser reactors, can have multiple feed inlet nozzles. The solid biomass material and the fluid hydrocarbon co-feed can therefore be processed in the catalytic cracking reactor even if both components are not miscible by feeding each component through a separate feed inlet nozzle.
In another embodiment the fluid hydrocarbon co-feed and the solid biomass material are mixed together prior to entry into a catalytic cracking reactor to provide a feed mixture comprising the fluid hydrocarbon co-feed and the solid biomass material. When mixing the fluid hydrocarbon co-feed and solid biomass material, the solid biomass material is preferably a torrefied and micronized biomass material as described herein before. The fluid hydrocarbon co-feed and the, optionally micronized and/or torrefied, solid biomass material may be mixed in any manner known to a skilled person to be suitable for mixing a viscous liquid and a solid. Preferably the fluid hydrocarbon co-feed and the, optionally micronized and/or torrefied, solid biomass material are mixed by means of shaking, stirring and/or extruding.
The feed mixture may be prepared just before entry to a catalytic cracking reactor or it may optionally be held in a stirred feed vessel before being forwarded to a catalytic cracking reactor.
Subsequently the solid biomass material and the fluid hydrocarbon co-feed are contacted with the catalytic cracking catalyst in a catalytic cracking reactor.
As indicated above, the catalytic cracking reactor is preferably a riser reactor. Preferably such a riser reactor is a riser reactor suitable for fluidized catalytic cracking. More preferably such a riser reactor is part of a catalytic cracking unit, more preferably of a fluidized catalytic cracking (FCC) unit.
In one preferred embodiment, a suspension of solid biomass material suspended in a fluid hydrocarbon feed is supplied to a riser reactor. Preferences for the fluid hydrocarbon feed are as described herein above.
In another preferred embodiment, the catalytic cracking reactor is a riser reactor and the solid biomass material is supplied to the riser reactor at a location downstream of a location where a fluid hydrocarbon feed is supplied to the riser reactor. Without wishing to be bound by any kind of theory it is believed that by allowing the fluid hydrocarbon co-feed to contact the catalytic cracking catalyst first, hydrogen may be generated. The availability of this hydrogen may assist in the reduction of coke formation when the solid biomass material is contacted with the catalytic cracking catalyst more downstream in the riser reactor.
In another preferred embodiment, the catalytic cracking reactor is a riser reactor and the solid biomass material is supplied to the riser reactor at a location upstream of a location where a fluid hydrocarbon feed is supplied to the riser reactor. Without wishing to be bound by any kind of theory it is believed that this allows the solid biomass material to be contacted with the catalytic cracking catalyst first; allowing the solid biomass material to be converted into an intermediate oil product and allowing this intermediate oil product to be at least partly and preferably wholly vaporized before the catalytic cracking catalyst is quenched by addition of a fluid hydrocarbon feed. In addition supplying the solid biomass material upstream of the fluid hydrocarbon feed may lead to in-situ water production in the upstream part of the riser reactor, leading to lower partial hydrocarbon pressures in the upstream part of the riser reactor and higher olefins yields. Further supplying the solid biomass material upstream of the fluid hydrocarbon feed allows for longer residence times for the solid biomass material, making it possible to use a solid biomass material with a particle size distribution having a particle size of equal to or more than 2000 micrometer.
In a still further embodiment, a suspension of solid biomass material suspended in a first fluid hydrocarbon feed is supplied to the riser reactor at a first location and a second fluid hydrocarbon feed is supplied to the riser reactor at a second location downstream of the first location. Preferences for the first and second fluid hydrocarbon feed are as described for the fluid hydrocarbon feed herein above.
By a riser reactor is herein understood an elongated, preferably essentially tube-shaped, reactor suitable for carrying out catalytic cracking reactions. Suitably a fluidized catalytic cracking catalyst flows in the riser reactor from the upstream end to the downstream end of the reactor. The elongated, preferably essentially tube-shaped, reactor is preferably oriented in an essentially vertical manner. Preferably a fluidized catalytic cracking catalyst flows from the bottom of the riser reactor upwards to the top of the riser reactor.
Examples of suitable riser reactors are described in the Handbook titled “Fluid Catalytic Cracking technology and operations”, by Joseph W. Wilson, published by PennWell Publishing Company (1997), chapter 3, especially pages 101 to 112, herein incorporated by reference.
For example, the riser reactor may be a so-called internal riser reactor or a so-called external riser reactor as described therein.
By an internal riser reactor is herein preferably understood an essentially vertical, preferably essentially tube-shaped, reactor, that may have an essentially vertical upstream end located outside a vessel and an essentially vertical downstream end located inside the vessel. The downstream end of the internal riser reactor that is located inside the vessel preferably comprises equal to or more than 30%, more preferably equal to or more than 40%, still more preferably equal to or more than 50% and most preferably equal to or more than 70% of the total length of the riser reactor. The vessel is suitably a reaction vessel suitable for catalytic cracking reactions and/or a vessel that comprises one or more cyclone separators and/or swirl tubes. The internal riser reactor is especially advantageous because in the process according to the invention, the solid biomass material may be converted into an intermediate oil product (also sometimes referred to as pyrolysis oil). Without wishing to be bound to any kind of theory it is believed that this intermediate oil product or pyrolysis oil may be more prone to polymerization than conventional oils due to oxygen-containing hydrocarbons and/or olefins that may be present in the intermediate oil product. By reducing polymerization of the olefins formed, also the overall olefin yield may be increased. In addition the intermediate oil product may be more corrosive than conventional oils due to oxygen-containing hydrocarbons that may be present. Further the internal riser reactor may be less sensitive to erosion by any unconverted particles of solid biomass material. The use of an internal riser reactor allows one to reduce the risk of plugging due to polymerization and/or to reduce the risk of corrosion and/or erosion, thereby increasing safety and hardware integrity.
By an external riser reactor is herein preferably understood a riser reactor that is located outside a vessel. The external riser reactor can suitably be connected via a so-called crossover to a vessel. Preferably the external riser reactor comprises a, preferably essentially vertical, riser reactor pipe. Such a riser reactor pipe is located outside a vessel. The riser reactor pipe may suitably be connected via a, preferably essentially horizontal, downstream crossover pipe to a vessel. The downstream crossover pipe preferably has a direction essentially transverse to the direction of the riser reactor pipe. The vessel may suitably be a reaction vessel suitable for catalytic cracking reactions and/or a vessel that comprises one or more cyclone separators and/or swirl separators.
When an external riser reactor is used, it may be advantageous to use an external riser reactor with a curve or low velocity zone at its upper end as for example illustrated in the Handbook titled “Fluid Catalytic Cracking technology and operations”, by Joseph W. Wilson, published by PennWell Publishing Company (1997), chapter 3, FIG. 3-7, herein incorporated by reference. The curve and/or low velocity zone may for example connect the riser reactor pipe and the so-called crossover pipe.
By a low velocity zone is herein preferably understood a zone or an area within the external riser reactor where the velocity of the, preferably fluidized, catalytic cracking catalyst shows a minimum. The low velocity zone may for example comprise an accumulation space located at the most downstream end of the upstream riser reactor pipe as described above, extending such riser reactor pipe beyond the connection with the crossover pipe. An example of a low velocity zone is the so-called “Blind Tee”.
It has been advantageously found that a part of the catalytic cracking catalyst may deposit in the curve or low velocity zone, thereby forming a protective layer against corrosion and/or erosion by the catalytic cracking catalyst and/or any residual solid particles and against corrosion by any oxygen-containing hydrocarbons.
In a preferred embodiment the solid biomass material is supplied to the riser reactor in the most upstream half, more preferably in the most upstream quarter, and even more preferably at the most upstream tenth of the riser reactor. Most preferably solid biomass material is supplied to the riser reactor at the bottom of this reactor. Addition of the solid biomass material in the upstream part, preferably the bottom, of the reactor may advantageously result in in-situ water formation at the upstream part, preferably the bottom, of the reactor. The in-situ water formation may lower the hydrocarbon partial pressure and reduce second order hydrogen transfer reactions, thereby resulting in higher olefin yields. Preferably the hydrocarbon partial pressure is lowered to a pressure in the range from 0.7 to 2.8 bar absolute (0.07 to 0.28 MegaPascal), more preferably 1.2 to 2.8 bar absolute (0.12 to 0.28 MegaPascal).
It may be advantageous to also add a lift gas at the bottom of the riser reactor. Examples of such a liftgas include steam, vaporized oil and/or oil fractions, and mixtures thereof. Steam is most preferred as a lift gas from a practical perspective. However, the use of a vaporized oil and/or oil fraction (preferably vaporized liquefied petroleum gas, gasoline, diesel, kerosene or naphtha) as a liftgas may have the advantage that the liftgas can simultaneously act as a hydrogen donor and may prevent or reduce coke formation. In an especially preferred embodiment both steam as well as vaporized oil and/or a vaporized oil fraction (preferably liquefied petroleum gas, vaporized gasoline, diesel, kerosene or naphtha) are used as a liftgas. Most preferably the liftgas consists of steam.
If the solid biomass material is supplied at the bottom of the riser reactor, is may optionally be mixed with such a lift gas before entry in the riser reactor. If the solid biomass material is not mixed with the liftgas prior to entry into the riser reactor it may be fed simultaneously with the liftgas (at one and the same location) to the riser reactor, and optionally mixed upon entry of the riser reactor; or it may be fed separately from any liftgas (at different locations) to the riser reactor.
When both the solid biomass material and the liftgas are introduced into the bottom of the riser reactor, the liftgas-to-solid biomass material weight ratio is preferably in the range from equal to or more than 0.01:1, more preferably equal to or more than 0.05:1 to equal to or less than 5:1, more preferably equal to or less than 1.5:1.
When the solid biomass material is introduced at the bottom of the riser reactor, it can be advantageous to increase the residence time of the solid biomass material at that part of the riser reactor by increasing the diameter of the riser reactor at the bottom. Hence in a preferred embodiment the riser reactor comprises a riser reactor pipe and a bottom section, which bottom section has a larger diameter than the riser reactor pipe, and wherein the solid biomass material is supplied to the riser reactor in the bottom section. This may advantageously lead to an improved conversion of the solid biomass material and less unconverted solid biomass material particles.
Where applicable a diameter is herein preferably understood to refer to the inner diameter, as for example the inner (i.e. internal) diameter of the bottom section or riser reactor pipe. Preferably the maximum inner diameter of the bottom section of the riser reactor is larger than the maximum inner diameter of the riser reactor pipe.
The bottom section having the larger diameter may for example have the form of a lift pot. The bottom section having the larger diameter is therefore also herein referred to as liftpot or enlarged bottom section.
Such a enlarged bottom section preferably has a diameter larger than the diameter of the riser reactor pipe, more preferably a diameter in the range from equal to or more than 0.4 to equal to or less than 5 meters, most preferably a diameter in the range from equal to or more than 1 to equal to or less than 2 meters. The height of the enlarged bottom section or liftpot preferably lies in the range from equal to or more than 1 meter to equal to or less than 5 meter.
In a further preferred embodiment the riser reactor may have a diameter that increases in a downstream direction to allow for the increasing gas volume generated during the conversion of the solid biomass material. The increase of diameter may be intermittent, resulting in two or more sections of the riser reactor having a fixed diameter, wherein each preceding section has a smaller diameter than the subsequent section, when going in a downstream direction; the increase of diameter may be gradual, resulting in a gradual increase of the riser reactor diameter in a downstream direction; or the increase in diameter may be a mixture of gradual and intermittent increases.
The length of the riser reactor may vary widely. For practical purposes the riser reactor preferably has a length in the range from equal to or more than 10 meters, more preferably equal to or more than 15 meters and most preferably equal to or more than 20 meters, to equal to or less than 65 meters, more preferably equal to or less than 55 meters and most preferably equal to or less than 45 meters.
Preferably the temperature in the catalytic cracking reactor (preferably riser reactor) ranges from equal to or more than 450° C., more preferably from equal to or more than 480° C., to equal to or less than 800° C., more preferably equal to or less than 750° C.
Preferably the temperature at the location where the solid biomass material is supplied lies in the range from equal to or more than 500° C., more preferably equal to or more than 550° C., and most preferably equal to or more than 600° C., to equal to or less than 800° C., more preferably equal to or less than 750° C.
In certain embodiments it can be advantageous to supply the solid biomass material to a location in a riser reactor where the temperature is slightly higher, for example where the temperature lies in the range from equal to or more than 700° C., more preferably equal to or more than 720° C., even more preferably equal to or more than 732° C. to equal to or less than 800° C., more preferably equal to or less than 750° C. Without wishing to be bound by any kind of theory, it is believed this may lead to a quicker conversion of the solid biomass material into the intermediate oil product.
Preferably the pressure in the catalytic cracking reactor (preferably riser reactor) ranges from equal to or more than 0.5 bar absolute to equal to or less than 10 bar absolute (0.05 MegaPascal to 1.0 MegaPascal), more preferably from equal to or more than 1.0 bar absolute to equal to or less than 6 bar absolute (0.1 MegaPascal to 0.6 MegaPascal).
When a riser reactor is used, preferably the total average residence time of the solid biomass material lies in the range from equal to or more than 1 second, more preferably equal to or more than 1.5 seconds and even more preferably equal to or more than 2 seconds to equal to or less than 10 seconds, preferably equal to or less than 5 seconds and more preferably equal to or less than 4 seconds.
Residence time as referred to in this patent application is based on the vapour residence at outlet conditions, that is, residence time includes not only the residence time of a specified feed (such as the solid biomass material) but also the residence time of its conversion products.
When the solid biomass material has a particle size distribution with a a mean particle size of equal to or more than 2000 micrometer, preferably a particle size in the range from 100 micrometer to 1000 micron, the total average residence time of the solid biomass material most preferably lies in the range from equal to or more than 1 seconds, preferably to equal to or less than 2.5 seconds.
When the solid biomass material has a mean particle size in the range from 30 micrometer to 100 micrometer the total average residence time of the solid biomass material most preferably lies in the range from equal to or more than 0.1 to equal to or less than 1 seconds.
The weight ratio of catalyst to feed (that is the total feed of solid biomass material and the fluid hydrocarbon feed)—herein also referred to as catalyst:feed ratio—preferably lies in the range from equal to or more than 1:1, more preferably from equal to or more than 2:1 and most preferably from equal to or more than 3:1 to equal to or less than 150:1, more preferably to equal to or less than 100:1, most preferably to equal to or less than 50:1.
The weight ratio of catalyst to solid biomass material (catalyst:solid biomass material ratio) at the location where the solid biomass material is supplied to the riser reactor preferably lies in the range from equal to or more than 1:1, more preferably from equal to or more than 2:1 and most preferably from equal to or more than 3:1 to equal to or less than 150:1, more preferably to equal to or less than 100:1, even more preferably to equal to or less than 50:1, most preferably to equal to or less than 20:1.
If the fluid hydrocarbon feed is introduced to the riser reactor downstream of the solid biomass material, the fluid hydrocarbon feed may preferably be introduced to the catalytic cracking reactor at a location where the solid biomass material already had a residence time in the range from equal to or more than 0.01 seconds, more preferably from equal to or more than 0.05 seconds, and most preferably from equal to or more than 0.1 seconds to equal to or less than 2 seconds, more preferably to equal to or less than 1 seconds, and most preferably to equal to or less than 0.5 seconds.
In a preferred embodiment the ratio between the total residence time for the solid biomass material to the total residence time for the fluid hydrocarbon feed (residence solid biomass material: residence hydrocarbon ratio) lies in the range from equal to or more than 1.01:1, more preferably from equal to or more than 1.1:1 to equal to or less than 3:1, more preferably to equal to or less than 2:1.
Preferably the temperature at the location in a riser reactor where the fluid hydrocarbon feed is supplied ranges from equal to or more than 450° C., more preferably from equal to or more than 480° C., to equal to or less than 650° C., more preferably to equal to or less than 600° C. Without wishing to be bound by any kind of theory, it is believe that the addition of the fluid hydrocarbon feed may quench the catalytic cracking catalyst and may therefore lead to a lower temperature at the location where it is added to the riser reactor.
Hence, preferably the solid biomass material is introduced to a riser reactor at a location with temperature T1 and the fluid hydrocarbon feed is introduced to the riser reactor at a location with temperature T2 and temperature T1 is higher than temperature T2. Preferably both T1 and T2 are equal to or more than 400° C., more preferably equal to or more than 450° C.
The catalytic cracking catalyst can be any catalyst known to the skilled person to be suitable for use in a cracking process. Preferably, the catalytic cracking catalyst comprises a zeolitic component. In addition, the catalytic cracking catalyst can contain an amorphous binder compound and/or a filler. Examples of the amorphous binder component include silica, alumina, titania, zirconia and magnesium oxide, or combinations of two or more of them. Examples of fillers include clays (such as kaolin).
The zeolite is preferably a large pore zeolite. The large pore zeolite includes a zeolite comprising a porous, crystalline aluminosilicate structure having a porous internal cell structure on which the major axis of the pores is in the range of 0.62 nanometer to 0.8 nanometer. The axes of zeolites are depicted in the ‘Atlas of Zeolite Structure Types’, of W. M. Meier, D. H. Olson, and Ch. Baerlocher, Fourth Revised Edition 1996, Elsevier, ISBN 0-444-10015-6. Examples of such large pore zeolites include FAU or faujasite, preferably synthetic faujasite, for example, zeolite Y or X, ultra-stable zeolite Y (USY), Rare Earth zeolite Y(═REY) and Rare Earth USY (REUSY). According to the present invention USY is preferably used as the large pore zeolite.
The catalytic cracking catalyst can also comprise a medium pore zeolite. The medium pore zeolite that can be used according to the present invention is a zeolite comprising a porous, crystalline aluminosilicate structure having a porous internal cell structure on which the major axis of the pores is in the range of 0.45 nanometer to 0.62 nanometer. Examples of such medium pore zeolites are of the MFI structural type, for example, ZSM-5; the MTW type, for example, ZSM-12; the TON structural type, for example, theta one; and the FER structural type, for example, ferrierite. According to the present invention, ZSM-5 is preferably used as the medium pore zeolite.
According to another embodiment, a blend of large pore and medium pore zeolites may be used. The ratio of the large pore zeolite to the medium pore size zeolite in the cracking catalyst is preferably in the range of 99:1 to 70:30, more preferably in the range of 98:2 to 85:15.
The total amount of the large pore size zeolite and/or medium pore zeolite that is present in the cracking catalyst is preferably in the range of 5 wt % to 40 wt %, more preferably in the range of 10 wt % to 30 wt %, and even more preferably in the range of 10 wt % to 25 wt % relative to the total mass of the catalytic cracking catalyst.
Preferably the catalytic cracking catalyst is contacted in a cocurrent flow configuration with a cocurrent flow of the, preferably solid, biomass material and optionally fluid hydrocarbon feed.
Catalytic cracking of a biomass material with a catalytic cracking catalyst as described herein is preferably carried out in a catalytic cracking unit, preferably a fluidized catalytic cracking unit.
In a preferred embodiment step a) comprises a catalytic cracking process comprising:
a catalytic cracking step comprising contacting the solid biomass material and the fluid hydrocarbon feed with a catalytic cracking catalyst at a temperature of more than 400° C. in a catalytic cracking reactor to produce one or more cracked products and a spent catalytic cracking catalyst;
a separation step comprising separating the one or more cracked products from the spent catalytic cracking catalyst;
a regeneration step comprising regenerating spent catalytic cracking catalyst to produce a regenerated catalytic cracking catalyst, heat and carbon dioxide; and
a recycle step comprising recycling the regenerated catalytic cracking catalyst to the catalytic cracking step.
The catalytic cracking step is preferably carried out as described herein before.
The separation step is preferably carried out with the help of one or more cyclone separators and/or one or more swirl tubes. Suitable ways of carrying out the separation step are for example described in the Handbook titled “Fluid Catalytic Cracking; Design, Operation, and Troubleshooting of FCC Facilities” by Reza Sadeghbeigi, published by Gulf Publishing Company, Houston Tex. (1995), especially pages 219-223 and the Handbook “Fluid Catalytic Cracking technology and operations”, by Joseph W. Wilson, published by PennWell Publishing Company (1997), chapter 3, especially pages 104-120, and chapter 6, especially pages 186 to 194, herein incorporated by reference. The cyclone separators are preferably operated at a velocity in the range from 18 to 80 meters/second, more preferably at a velocity in the range from 25 to 55 meters/second.
In addition the separation step may further comprise a stripping step. In such a stripping step the spent catalyst may be stripped to recover the products absorbed on the spent catalyst before the regeneration step. These products may be recycled and added to the cracked product stream obtained from the catalytic cracking step.
The regeneration step preferably comprises contacting the spent catalytic cracking catalyst with an oxygen containing gas in a regenerator at a temperature of equal to or more than 550° C. to produce a regenerated catalytic cracking catalyst, heat and carbon dioxide. During the regeneration coke, that can be deposited on the catalyst as a result of the catalytic cracking reaction, is burned off to restore the catalyst activity.
The oxygen containing gas may be any oxygen containing gas known to the skilled person to be suitable for use in a regenerator. For example the oxygen containing gas may be air or oxygen-enriched air. By oxygen enriched air is herein understood air comprising more than 21 vol. % oxygen (O2), more preferably air comprising equal to or more than 22 vol. % oxygen, based on the total volume of air.
The heat produced in the exothermic regeneration step is preferably employed to provide energy for the endothermic catalytic cracking step. In addition the heat produced can be used to heat water and/or generate steam. The steam may be used elsewhere in the refinery, for example as a liftgas in the riser reactor.
Preferably the spent catalytic cracking catalyst is regenerated at a temperature in the range from equal to or more than 575° C., more preferably from equal to or more than 600° C., to equal to or less than 950° C., more preferably to equal to or less than 850° C. Preferably the spent catalytic cracking catalyst is regenerated at a pressure in the range from equal to or more than 0.5 bar absolute to equal to or less than 10 bar absolute (0.05 MegaPascal to 1.0 MegaPascal), more preferably from equal to or more than 1.0 bar absolute to equal to or less than 6 bar absolute (0.1 MegaPascal to 0.6 MegaPascal).
The regenerated catalytic cracking catalyst can be recycled to the catalytic cracking step. In a preferred embodiment a side stream of make-up catalyst is added to the recycle stream to make-up for loss of catalyst in the reaction zone and regenerator.
In step (a) of the process according to the invention one or more cracked products are produced. At least one of the one or more cracked products is subsequently fractionated in step b) to obtain one or more product fractions.
As indicated herein, the one or more cracked products may contain one or more oxygen-containing-hydrocarbons. Examples of such oxygen-containing-hydrocarbons include ethers, esters, ketones, acids and alcohols. In specific the one or more cracked products may contain phenols.
Preferably the one or more cracked product(s) produced in step (a) and fractionated in step (b) have an elemental oxygen content in the range from equal to or more than 0.01 wt %, more preferably equal to or more than 0.1 wt %, even more preferably equal to or more than 0.2 wt % and most preferably equal to or more than 0.3 wt % oxygen, to equal to or less than 10 wt %, more preferably equal to or less than 5 wt %, and most preferably equal to or less than 1 wt %, based on total weight of dry matter.
Fractionation may be carried out in any manner known to the skilled person in the art to be suitable for fractionation of products from a catalytic cracking reactor. For example the fractionation may be carried out as described in the Handbook titled “Fluid Catalytic Cracking technology and operations”, by Joseph W. Wilson, published by PennWell Publishing Company (1997), chapter 8, especially pages 223 to 235, herein incorporated by reference.
The one or more cracked products are preferably obtained as gaseous cracked products from step (a). These gaseous cracked products can subsequently be separated into various gas and liquid products in one or more fractionation units.
It has been found that the use of a solid biomass material as part of the feed in step (a) may lead to additional coke formation in conduits that transport the one or more cracked products from the catalytic cracking reactor. It has therefore been found advantageous to use insulated pipelines in step (b) to transfer the one or more cracked products from the catalytic cracking reactor into one or more subsequent fractionation units. Most preferably so called cold-wall pipelines are used which comprise an insulation on the inside pipeline surface.
Preferably a main fractionator is used to cool the gaseous cracked products obtained from step (a) and to condense any heavy liquid products. The main fractionator preferably comprises a distillation tower comprising a bottom section (sometimes referred to as flash zone) at the bottom of the tower; a heavy cycle oil (HCO) section, a light cycle oil (LCO) section and a top section.
In the bottom section the cracked products are preferably cooled via contact with a circulating stream of fractionator bottoms product, sometimes also referred to as a bottoms pump-around. In addition to cooling gaseous cracked products, the circulating liquid fractionator bottoms product can advantageously also be used to wash out any residual solid biomass particles.
In order to deal with any residual solid biomass particles the bottom section is preferably also fitted with one or more baffle plates, grid packing and/or one or more solid biomass particle catchers. Residual solid biomass particles accumulated in these catchers may advantageously be recycled to step a). The product obtained from the bottom section at the bottom of the tower is sometimes also referred to as slurry oil. By slurry oil is herein preferably understood a fraction of the cracked products of which at least 80 wt %, more preferably at least 90 wt % boils at or above 425° C. (at 0.1 MegaPascal). The slurry oil may still contain solid biomass particles that were not converted in the catalytic cracking reactor. Such solid biomass particles can be separated from the slurry oil by means of settling, filtration and/or electrostatic filtration and advantageously recycled to step a).
In the heavy cycle oil (HCO) section so called heavy cycle oil may be withdrawn from the distillation tower. By heavy cycle oil is herein preferably understood a fraction of the cracked products of which at least 80 wt %, more preferably at least 90 wt % boils in the range from equal to or more than 370° C. to less 425° C. (at 0.1 MegaPascal). In a preferred embodiment, at least part of this heavy cycle oil is advantageously recycled and used as a fluid hydrocarbon co-feed in step a).
When cooling the one or more cracked products to retrieve slurry oil in de bottom section and/or heavy cycle oil in the HCO section, the heat of the one or more cracked products can be retrieved and advantageously used to preheat the feed in step a). For example one or more of the feed streams for step a) can be used to cool the circulating liquid fractionator bottoms product.
In addition, at least part of the heavy cycle oil and/or at least part of the slurry oil may be used as a fuel to provide the heat for an optional torrefaction step as described herein before.
In the light cycle oil (LCO) section so called light cycle oil may be withdrawn from the distillation tower. By light cycle oil (LCO) is herein preferably understood a fraction of the cracked products of which at least 80 wt %, more preferably at least 90 wt % boils in the range from equal to or more than 221° C. to less than 370° C. (at 0.1 MegaPascal). This light cycle oil or parts thereof can advantageously be hydrodeoxygenated in step (c) to prepare the one or more hydrodeoxygenated product(s) as described in more detail below. Alternatively, at least part of the light cycle oil may also be withdrawn and directly used as a biofuel component and/or a biochemical component.
In the top section of the distillation tower, naphtha products and so-called dry gas can be withdrawn. By naphtha products is herein preferably understood a fraction of the cracked products of which at least 80 wt %, more preferably at least 90 wt % boils in the range from equal to or more than 30° C. to less than 221° C. (at 0.1 MegaPascal).
By dry gas is herein preferably understood a fraction consisting of compounds boiling at or below the boiling point of ethane. The dry gas may comprise for example methane, ethane, ethene, carbon-monoxide, carbon-dioxide, hydrogen and nitrogen. The naphtha products may comprise fractions that may be useful as a biofuel component for gasoline compositions and/or diesel compositions. Preferably the dry gas is separated from the naphtha products by means of one or more gas/liquid separator(s) and/or one or more absorber(s). Subsequently the naphtha products may be debutanized and/or depentanized, if so desired, to remove compounds boiling at or below the boiling point of butane respectively below the boiling point of pentane. At least part of the, optionally debutanized and/or depentanized, naphtha products may be advantageously hydrodeoxygenated in step (c) to prepare one or more hydrodeoxygenated products as described in more detail below. Alternatively, at least part of the naphtha products may also be withdrawn and directly used as a biofuel component and/or a biochemical component.
In a further embodiment the, optionally debutanized and/or depentanized, naphtha products are forwarded to one or more further distillation column(s). Here the, optionally debutanized and/or depentanized, naphtha products may be split up into a light-light-cycle oil (LLCO, sometimes also referred to as heavycat-cracked gasoline (HCCG)); a cat-cracked gasoline (CCG, sometimes also referred to as heart cut CCG); and a light cat-cracked gasoline (LCCG, sometimes also referred to as cat-cracked tops). By light cat-cracked gasoline is herein preferably understood a fraction of the naphtha products of which at least 80 wt %, more preferably at least 90 wt % boils in the range from equal to or more than 35° C. to less than 125° C. (at 0.1 MegaPascal). If desired, each of the light-light cycle oil, heavy cat cracked gasoline and/or light cat-cracked gasoline may independently be hydrodeoxygenated as described herein below for step c).
As indicated above, fractionation of at least one of the one or more cracked products in step (b) results in one or more product fractions. At least one of the one or more product fractions obtained in step (b) is subsequently hydrodeoxygenated in step (c) to produce one or more hydrodeoxygenated products.
Examples of product fractions that can be obtained from step (b) and can be hydrodeoxygenated in step (c) include naphtha products, such as gasoline or diesel fractions; light cycle oils (LCO); heavy cycle oil (HCO); slurry oil; fractions thereof and/or mixtures thereof.
In a preferred embodiment, the one or more product fractions obtained in step (b) and subsequently hydrodeoxygenated in step (c) contain naphtha products; naphtha product fractions (such as for example heavy cat cracked gasoline); light cycle oils (LCO); LCO fractions; and/or mixtures thereof.
Preferably the one or more product fractions obtained in step (b) and subsequently hydrodeoxygenated in step (c) consists of a fraction of the cracked products of which at least 70 wt %, more preferably at least 80 wt %, most preferably at least 90 wt % boils in the range from equal to or more than 30° C. to less than 370° C. (at 0.1 MegaPascal). More preferably the one or more product fractions obtained in step (b) and subsequently hydrodeoxygenated in step (c) consists of a fraction of the cracked products of which at least 70 wt %, more preferably at least 80 wt %, most preferably at least 90 wt % boils in the range from equal to or more than 30° C. to less than 221° C. (at 0.1 MegaPascal).
The one or more product fractions obtained in step (b) and subsequently hydrodeoxygenated in step (c) may contain one or more oxygen-containing-hydrocarbons. Examples of such oxygen-containing-hydrocarbons include ethers, esters, ketones, acids and alcohols. In specific one or more product fractions may contain phenols and/or substituted phenols.
Preferably the one or more product fraction(s) produced in step (b) and hydrodeoxygenated in step (c) have an elemental oxygen content in the range from equal to or more than 0.01 wt %, more preferably equal to or more than 0.1 wt %, even more preferably equal to or more than 0.2 wt % and most preferably equal to or more than 0.3 wt % oxygen, to equal to or less than 20 wt %, more preferably equal to or less than 10 wt %, and most preferably equal to or less than 5 wt %, based on total weight of dry matter (i.e. essentially water-free basis).
By hydrodeoxygenation is herein understood reducing the concentration of oxygen-containing hydrocarbons in one or more product fraction(s) containing oxygen-containing hydrocarbons by contacting the one or more product fraction(s) with hydrogen in the presence of a hydrodeoxygenation catalyst. Oxygen-containing hydrocarbons that can be removed include acids, ethers, esters, ketones, aldehydes, alcohols (such as phenols) and other oxygen-containing compounds.
The hydrodeoxygenation preferably comprises contacting of the one or more product fractions with hydrogen in the presence of an hydrodeoxygenation catalyst at a temperature in the range from equal to or more than 200° C., preferably equal to or more than 250° C., to equal to or less than 450° C., preferably equal to or less than 400° C.; at a total pressure in the range of equal to or more than 10 bar absolute (1.0 MegaPascal) to equal to or less than 350 bar absolute (35 MegaPascal); and at a partial hydrogen pressure in the range of equal to or more than 2 bar absolute (0.2 MegaPascal) to equal to or less than 350 bar absolute (35 MegaPascal).
When the product fraction is a light cycle oil (LCO) or a fraction thereof, hydrodeoxygenation is more preferably carried out at a total pressure in the range of equal to or more than 30 bar absolute (3.0 MegaPascal) and most preferably equal to or more than 50 bar absolute (5.0 MegaPascal) to equal to or less than 350 bar absolute (35 MegaPascal), more preferably equal to or less than 300 bar absolute (30 MegaPascal); and at a partial hydrogen pressure in the range of equal to or more than 20 bar absolute (2.0 MegaPascal) and most preferably equal to or more than 40 bar (4.0 MegaPascal) absolute to equal to or less than 350 bar absolute (35 MegaPascal), more preferably equal to or less than 300 bar absolute (30 MegaPascal).
When the product fraction is naphtha or a fraction thereof, hydrodeoxygenation is more preferably carried out at a total pressure in the range of equal to or more than 10 bar absolute (1.0 MegaPascal) and most preferably equal to or more than 20 bar absolute (2.0 MegaPascal) to equal to or less than 100 bar absolute (10 MegaPascal), more preferably equal to or less than 60 bar absolute (6.0 MegaPascal); and at a partial hydrogen pressure in the range of equal to or more than 5 bar absolute (0.5 MegaPascal) and most preferably equal to or more than 10 bar absolute (1.0 MegaPascal) to equal to or less than 100 bar absolute (10 MegaPascal), more preferably equal to or less than 60 bar absolute (6.0 MegaPascal).
The hydrodeoxygenation catalyst can be any type of hydrodeoxygenation catalyst known by the person skilled in the art to be suitable for this purpose.
The hydrodeoxygenation catalyst preferably comprises one or more hydrodeoxygenation metal(s), preferably supported on a catalyst support. The catalyst support is preferably inert as a hydrodeoxygenation catalyst at the hydrodeoxygenation conditions.
The one or more hydrodeoxygenation metal(s) are preferably chosen from Group VIII and/or Group VIB of the Periodic Table of Elements. The hydrodeoxygenation metal may for example be present as a mixture, alloy or organometallic compound.
Preferably the one or more hydrodeoxygenation metal(s) is chosen from the group consisting of Nickel (Ni), Chromium (Cr), Molybdenum (Mo), Tungsten (W), Cobalt (Co), Platinum (Pt), Palladium (Pd), Rhodium (Rh), Ruthenium (Ru), Iridium (Ir), Osmium (Os), Copper (Cu), iron (Fe), Zink (Zn), Gallium (Ga), Indium (In), Vanadium (V) and mixtures thereof. The one or more metal(s) may be present in elementary form; in the form of alloys or mixtures; and/or in the form of oxides, sulfides or other metal-organic compounds.
Preferably the hydrodeoxygenation catalyst in step (c) is a catalyst comprising Tungsten, Ruthenium, Rhenium, Cobalt, Nickel, Copper, Molybdenum, alloys thereof and/or mixtures thereof.
Most preferably the hydrodeoxygenation catalyst is chosen from the group consisting of Rhodium-Cobalt catalysts, Nickel-Tungsten catalysts, Nickel-Copper catalysts, Cobalt-Molybdenum catalysts and Nickel-Molybdenum catalyst. Such Rhodium-Cobalt catalysts, Nickel-Tungsten catalysts, Nickel-Copper catalysts, Cobalt-Molybdenum catalysts or Nickel-Molybdenum catalysts may contain such metals supported on an inert catalyst support as described above.
If the hydrodeoxygenation catalyst comprises a catalyst support, such catalyst support may be shaped in the form of balls, rings or otherwise shaped extrudates. The catalyst support may comprise a refractory oxide or mixtures thereof, preferably alumina, amorphous silica-alumina, titania, silica, ceria, zirconia; or it may comprise an inert component such as carbon or silicon carbide. Preferred are for example ZrO2, CeO2, CeO2 and/or mixtures thereof. The catalyst support may further comprise a zeolitic compound such as for example zeolite Y, zeolite beta, ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-48, SAPO-11, SAPO-41, and ferrierite.
Examples of suitable catalysts include Rh/SiO2; RhCo/Al2O3; Rh/CoSiO3; RhCo/SiO2; Co/SiO2; Rh/ZrO2; Rh/CeO2; Ni/SiO2; Ni/Cr2O3; Ni/Al2O3; Ni/ZrO2; Ni—Cu/Al2O3; Ni—Cu/ZrO2; Ni—Cu/CeO2; Ni—Mo/Al2O3; Ni—Mo/ZrO2; Co—Mo/Al2O3 and Co—Mo/ZrO2. Preferably the catalyst is chosen from the group consisting of Rh/Al2O3, RhCo/Al2O3; Rh/ZrO2; Rh/CeO2; Ni/Cr2O3; Ni/Al2O3; Ni/ZrO2; Ni—Cu/Al2O3; NiW/Al2O3; Ni—Cu/ZrO2; Ru/C; Ni—Cu/CeO2; Ni—Mo/Al2O3; Ni—Mo/ZrO2; Co—Mo/Al2O3; Co—Mo/ZrO2 and/or mixtures thereof.
Most preferred are hydrodeoxygenation catalysts comprising Rhodium on alumina (Rh/Al2O3), Rhodium-Cobalt on alumina (RhCo/Al2O3), Nickel-Copper on alumina (NiCu/Al2O3), Nickel-Tungsten on alumina (NiW/Al2O3), Cobalt-Molybdenum on alumina (CoMo/Al2O3) or Nickel-Molybdenum on alumina (NiMo/Al2O3).
If the one or more product fractions also contain one or more sulphur-containing hydrocarbons it may be advantageous to use a sulphided hydrodeoxygenation catalyst. If the hydrodeoxygenation catalyst is sulphided the catalyst may be sulphided in-situ or ex-situ. Such in-situ or ex-situ sulphiding can be carried out in any manner known by the skilled person to be suitable for in-situ or ex-situ sulphiding. In the case of in-situ sulphiding, a sulfur source, usually hydrogen sulphide or a hydrogen sulphide precursor, is preferably supplied to the hydrodeoxygenation catalyst before operation of the process in a hydrodeoxygenation reactor. In addition it may be advantageous to add a small amount of hydrogen sulphide during operation of the hydrodeoxygenation process to keep the catalyst sufficiently sulphided.
In addition to the hydrodeoxygenation, step (c) may comprise further steps, if so desired or necessary. For example, if desired, step (c) may further comprise hydrodesulphurization, hydrodenitrogenation, hydrocracking and/or hydroisomerization of the one or more product fractions. Hydrodesulphurization may reduce the concentration of any sulphur-containing hydrocarbons. Hydrodenitrogenation may reduce the concentration of any nitrogen-containing hydrocarbons. Hydroisomerization may increase the concentration of branched hydrocarbons. Hydrocracking may further crack the product in smaller compounds.
Such hydrodesulphurization, hydrodenitrogenation, hydrocracking and/or hydroisomerization may be carried out before, after and/or simultaneously with the hydrodeoxygenation.
The hydrodeoxygenation can be carried out in any type of reactor known by the skilled person in the art to be suitable for a hydrodeoxygenation process. Preferably a fixed bed reactor, trickle flow reactor, ebullated bed reactor or fluidized bed reactor is used. In a preferred embodiment a weight hourly space velocity is used that is equal to or more than 0.2 and equal to or less than 4.0 kg/litre hour.
In step (c) one or more hydrodeoxygenated product(s) may be obtained. These one or more hydrodeoxygenated product(s) can be used as biofuel component(s) and/or a biochemical component(s). By a biofuel component is herein understood a component that can be useful in the preparation of a biofuel. By a biochemical component is herein understood a component that can be useful in the preparation of a biochemical.
Preferably the one or more hydrodeoxygenated product(s) has an elemental oxygen content of equal to or less than 0.03 wt % (300 ppmw), more preferably equal to or less than 0.01 wt %. Most preferably the one or more hydrodeoxygenated product(s) is essentially free of oxygen-containing hydrocarbons and/or contains essentially no elemental oxygen.
In a preferred embodiment the one or more hydrodeoxygenated product(s) produced in step (c) can be blended with one or more other components to produce a biofuel and/or a biochemical. Examples of one or more other components with which the one or more hydrodeoxygenated product(s) may be blended include anti-oxidants, corrosion inhibitors, ashless detergents, dehazers, dyes, lubricity improvers and/or mineral fuel components.
Alternatively the one or more hydrodeoxygenated product(s) can be used in the preparation of a biofuel component and/or a biochemical component. In such a case the biofuel component and/or biochemical component prepared from the one or more hydrodeoxygenated product may be subsequently blended with one or more other components (as listed above) to prepare a biofuel and/or a biochemical.
By a biofuel respectively a biochemical is herein understood a fuel or a chemical that is at least party derived from a renewable energy source.
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The invention is further illustrated by the following, non-limiting examples.
The invention will now be further illustrated by means of the following non-limiting examples. In the experiments below, elemental carbon and hydrogen analysis has been carried out by ASTM method D5291 unless indicated otherwise; elemental nitrogen analysis has been carried out by ASTM method D5762 (for nitrogen >100 ppmw) and ASTM 4629 (for nitrogen less than 100 ppm) on an Antek 9000 apparatus, unless indicated otherwise; elemental oxygen analysis has been carried out on a Eurovector EA3000 apparatus (commercially available from Eurovector) or an Vario Cube apparatus of Elementar GmbH, unless indicated otherwise; elemental sulphur analysis is carried out by ASTM D5453 combustion by UV-fluorescence detection on an Antek 9000 apparatus, unless indicated otherwise. Gas Chromatography and oxygen content of the feed were determined with ASTM D5599-95 unless indicated otherwise.
Chips of poplar wood were torrefied at 250° C. for 6 hours. These were finely milled using a Retch PM 400 ball mill for 4 hours and at 400 rpm to produce a micronized torrefied wood material. The milled torrefied poplar wood had an apparent Bulk Density of 0.42 g/mL and a mean particle size distribution of 36 micrometer (as measured at a Horiba LA950, Laser Scattering Particle Size Distribution Analyzer). The milled torrefied poplar wood was dried during 1 day at 105° C. Subsequently the milled torrefied poplar wood (MTPW) was added to a fluid hydrocarbon feed (HF) in weight ratios as indicated below:
The above combinations of milled torrefied poplar wood and fluid hydrocarbon feed were shaken and stirred for 1 hour at room temperature and mixed/shredded for 1 hour with an Inline Ultra Turrax extruder at 80° C. to give a pumpable mixture. The boiling range distribution of the fluid hydrocarbon feed is provided in table 1. An elemental analysis of the milled torrefied poplar wood and of the fluid hydrocarbon feed is provided in table 2. The milled torrefied poplar wood was dried to remove water before analysis.
Mixtures of milled torrefied poplar wood and fluid hydrocarbon feed as listed under 1a, 1b, 1c and 1d were injected as feed mixtures from a stirred feed vessel at 60° C. into a fluidized bed of a MAT-5000 fluidized catalytic cracking unit.
Each mixture was tested in a separate run. Each run included 7 experiments with 7 catalyst to feed weight ratios, namely catalyst to feed (Cat/feed) weight ratios of 3, 4, 5, 6, 6.5, 7 and 8.
Each experiment was conducted as follows: 10 g of FCC equilibrium catalyst (i.e. the catalytic cracking catalyst) containing ultra stable zeolite Y with a composition as listed in table 3, was constantly fluidized with nitrogen. A precise and known amount of feed was injected, and subsequently flushed through a tube with nitrogen to the fluidized catalyst bed during a runtime of 1 minute. The fluidized catalyst bed was kept at 520° C. The liquid products, also referred to as total liquid product (TLP), were collected in glass receivers at minus 18-19° C. Subsequently the catalytic cracking catalyst was stripped with nitrogen. The gas produced during such stripping was weighted and analyzed online with a gas chromatograph (GC). Hereafter the catalytic cracking catalyst was regenerated in-situ at 650° C. in the presence of air. During such regeneration the coke was converted to CO2, which was quantified by on-line infrared measurement. After regeneration the reactor was cooled to the cracking temperature and a new injection was started. One cycle including all catalyst to oil ratios took approximately 16 hours.
The 7 experiments were used to generate extrapolated results for each catalyst to feed weight ratio and each conversion. From these extrapolated results the yields for the different products fractions are listed in table 4 were determined. The elemental analysis of the total liquid product is listed in table 5.
As illustrated in table 5, the TLP still contains elemental oxygen. By hydrodeoxygenation of the total liquid product listed in table 5, oxygen can be removed.
The biocarbon content of the Total Liquid Product obtained by catalytic cracking of the feed mixture containing 20 wt % milled torrefied poplar wood and 80 wt % Fluid hydrocarbon feed was measured by means of carbon isotope analysis according to ASTM 6866. Measurement showed that about 38 wt % of the elemental carbon (also referred to as bio-carbon) present in the milled torrefied poplar wood was present in the catalytically cracked total liquid product. It has been determined that the overall mass balances (torrefied wood and fluid hydrocarbon feed) for each and every experiment above are in the range 98-102 wt % showing that the torrefied wood feed is introduced in its totality into the reactor. The variation in mass balance is the normally observed experimental variation in FCC experiments in the used reactor.
A feed mixture containing 18 wt % milled torrefied poplar wood and 82 wt % fluid hydrocarbon feed with properties as listed in tables 1 and 2 was catalytically cracked at 520° C. with a catalyst to feed weight ratio of 3 in a MAT-5000 fluidized catalytic cracking unit and with a catalytic cracking catalyst as described in example 2 to produce a Total Liquid Product (TLP).
In an 500 ml autoclave (Ernst Haage), 10 grams of this Total Liquid Product was diluted with 150 ml dodecane and subsequently blended with 0.77 grams of a pre-sulphided hydrodeoxygenation catalyst containing 3.5 wt % Nickel and 15 wt % Molybdenum on alumina (DN3531, commercially obtainable from Criterion). The autoclave was pressurized with a gas mixture of 1% hydrogensulfide (H2S) in hydrogen (H2) to 60 bars (6.0 MegaPascal). Hereafter the autoclave was heated to 300° C. (at a rate of 25° C. per minute) resulting in a final pressure of 87 bar (8.7 MegaPascal). After processing the above in the autoclave for about 2 hours, the autoclave was cooled and depressurized. Samples were taken from the Total Liquid Product before and after hydrodeoxygenation for elemental analysis. The results of the elemental analysis are shown in table 6. Elemental analysis of the hydrodeoxygenated Total Liquid Product showed the essential absence of oxygen.
| Number | Date | Country | Kind |
|---|---|---|---|
| 11163542.1 | Apr 2011 | EP | regional |
