Treatment of BIO-oil feedstock

Abstract

A process is disclosed for catalytic deoxygenation of a feedstock of bio-oils derived from lignocellulosic biomass, comprising hydrotreating a slurry comprising the feedstock of bio-oils derived from lignocellulosic biomass and an unsuported molybdenum-based catalyst with hydrogen at a temperature of 270 to 410 °C, preferably 350 to 400 °C, more preferably 370 to 390 °C, and at a pressure of 30 to 200 bar.

Description

[0001] TREATMENT OF BIO-OIL FEEDSTOCK

[0002] FIELD OF THE INVENTION

[0003] The present invention relates to a process for catalytic deoxygenation of a feedstock of bio-oils derived from lignocellulosic biomass.

[0004] The present invention also relates to a composition obtainable by said process for catalytic deoxygenation of a feedstock of bio-oils derived from lignocellulosic biomass.

[0005] BACKGROUND OF THE INVENTION

[0006] Bio-oils produced by thermochemical liquefaction methods (thermal fast pyrolysis, catalytic fast pyrolysis, hydrothermal liquefaction) typically contain 15 - 40 wt-% of oxygen thus deviating significantly from hydrocarbon transportation fuels. Furthermore, these bio-oils have low thermal stability, low energy density and low miscibility in hydrocarbons, and typically they also have a high impurity content. State-of-the-art technology proposes to upgrade these bio-oils to transportation fuels with oxygen content close to zero by hydrotreatment using supported metal catalysts in fixed bed reactors. However, this process has not been yet commercialized due to problems related to fast catalyst deactivation caused by bio-oils’ instability (coke formation) and impurities in the oils (catalyst poisoning). Therefore, new methods are needed to upgrade bio-oils from liquefaction processes to advanced transportation fuels and green chemical intermediates.

[0007] Hydroprocessing of bio-oils is known in the art. Publications US20170022425A1 and US10920151B2 teach a staged fixed bed hydrotreatment approach in which a series of fixed bed reactors is applied to hydrotreat a bio-oil feedstock.

[0008] Publication W02023156742A1 discloses a method of producing a stabilized biomass oil, also called bio-oils produced by pyrolysis or hydrothermal liquefaction of biomass.

[0009] Publication W02023175160A1 discloses a process for conversion of solid lignocellulosic raw materials to an organic liquefaction product.

[0010] Publication W02023175155A1 also discloses a process for conversion of solid lignocellulosic raw materials to an organic liquefaction product, as well as a sulfided molybdenum catalyst for use in said process, and methods to produce said catalyst.

[0011] Publication W02023187732A1 discloses a directprocess for hydroconversion of biomass. The process comprises separately feeding a solid biomassfeedstock and a liquid feedstock to a slurry hydroconversion reactor. The slurry hydroconversion reactor comprises a slurry hydroconversion catalyst. The solid biomass feedstock and the liquid feedstock are contacted with the slurry hydroconversion catalyst for a sufficient time under hydroconversion process conditions in the presence of hydrogen to convert the solid biomass feedstock and the liquid feedstock to hydroconversion product. The catalyst comprises an unsupported catalyst selected from molybdenum sulfide, iron sulfide, nickel sulfide, zinc sulfide, iron zinc, or a combination thereof. The catalyst has an average particle size of from 1 to 100 microns, e.g. from 2 to 10 microns.

[0012] Publication EP4100491A1 (W02021 / 156436) discloses a process of producing a hydrocracking product in a slurry hydrocracking reactor, in which process

[0013] - a pyrolysis oil, a hydrocarbon feedstock, and a hydrocracking catalyst is provided; - the pyrolysis oil is combined with the hydrocarbon feedstock and the hydrocracking catalyst, the pyrolysis oil being maintained at a temperature of less than 100 °C until the pyrolysis oil contacts both the hydrocarbon feedstock and the hydrocracking catalyst;

[0014] - the hydrocarbon feedstock and the pyrolysis oil are hydrocracked in the slurry hydrocracking reactor in the presence of the hydrocracking catalyst and hydrogen gas.

[0015] BRIEF DESCRIPTION OF THE INVENTION

[0016] The scope of protection sought for various example embodiments is set out by the independent claims. The example embodiments and features, if any, described in this specification that do not fall under the scope of the independent claims are to be interpreted as examples useful for understanding various embodiments.

[0017] According to an aspect, there is provided a process for catalytic deoxygenation of a feedstock of bio-oils derived from lignocellulosic biomass, comprising:

[0018] hydrotreating a slurry comprising the feedstock of bio-oils derived from lignocellulosic biomass and an unsupported molybdenum-based catalyst with hydrogen ata temperature of 270 to 410 °C, preferably 350 to 400 °C, more preferably 370 to 390 °C, and at a pressure of 30 to 200 bar.

[0019] According to another aspect, there is provided a composition of deoxygenated bio-oils obtainable by the process according to the invention.BRIEF DESCRIPTION OF THE DRAWINGS

[0020] In the following the invention will be described in greater detail by means of preferred embodiments with reference to the attached drawings, in which

[0021] Figure 1 is a schematic representation of a continuous slurry hydroprocessing reactor unit;

[0022] Figure 2 shows a particle size distribution of an exemplary catalyst. DETAILED DESCRIPTION OF THE INVENTION

[0023] The following embodiments are exemplifying. Although the specification may refer to “an”, “one”, or “some” embodiment(s) in several locations of the text, this does not necessarily mean that each reference is made to the same embodiment's), or that a particular feature only applies to a single embodiment. Single features of different embodiments may also be combined to provide other embodiments.

[0024] In an embodiment, a process is disclosed for catalytic deoxygenation of a feedstock of bio-oils derived from lignocellulosic biomass, the process comprising hydrotreating a slurry comprising the feedstock of bio-oils derived from lignocellulosic biomass and an unsupported molybdenum-based catalyst with hydrogen at a temperature of 270 to 410 °C, preferably 350 to 400 °C, more preferably 370 to 390 °C, and at a pressure of 30 to 200 bar.

[0025] In an embodiment the feedstock of bio-oils derived from lignocellulosic biomass may comprise bio-oils obtained by pyrolysis of lignocellulosic biomass, preferably fast pyrolysis of lignocellulosic biomass, or by liquefaction of lignocellulosic biomass. Fast pyrolysis may be pyrolysis with low residence time.

[0026] In further embodiment the unsupported molybdenum-based catalyst comprises a promoter metal (Me) and has the formula MoxMeySz, wherein x=l, y=0.3-1.5, z=3-5, preferably x=l, y=0.5-0.8, z=3-4, and Me is Co or Ni.

[0027] In another embodiment the concentration of the unsupported molybdenum-based catalyst in the slurry may be 100-10000 ppm by weight, preferably 200-5000 ppm by weight of molybdenum metal present in the catalyst.

[0028] In a further embodiment the solid content of the slurry may be less than 5 wt-%, preferably less than 1 wt-%, more preferably less than 0.1 wt-%.

[0029] In a further embodiment the unsupported molybdenum-based catalyst may have a particle size D 50 of 5 to 100 microns.

[0030] In a further embodiment the pressure may be 50-200 bar, preferably60-200 bar.

[0031] In a still further embodiment the hydrotreating may be performed in a slurry reactor.

[0032] In a further embodiment the process may comprise continuously replacing unsupported molybdenum-based catalyst spent during hydrotreating the slurry with fresh amounts of said unsupported molybdenum-based catalyst.

[0033] In a further embodiment, the process may further comprise a separation step, wherein a gas phase and an aqueous phase formed during hydrotreating the slurry may be separated from an organic liquid phase of the slurry and the gas phase may be optionally partly circulated back to the slurry.

[0034] In a further embodiment light and heavy organic phases may be separated from each other and the heavy organic phase containing catalyst may be optionally partly or fully circulated back to the slurry.

[0035] In a further embodiment of the process of the invention the spent catalyst may be removed from the slurry with the part of the heavy organic phase not circulated back to the slurry, and fresh catalyst is added to the slurry.

[0036] According to another aspect there is provided a composition of deoxygenated bio-oils obtainable by the process or any of its embodiments.

[0037] In an embodiment, the composition may have

[0038] - a kinematic viscosity of 35 mm / s2or less, preferably in the range of 4-25 mm / s2

[0039] - a density of 1.06 g / ml or less, preferably in the range of 1 to 0.9 g / ml - a dry oxygen content of less than 25 wt-%, preferably in the range of 10-25 wt-%.

[0040] In a slurry hydroprocessing process, catalyst may be continuously added and removed solving the problem of catalyst fast deactivation process in biooils upgrading by hydroprocessing. Unsupported Mo and promoted Mo catalysts in sulfide form may be suitable catalysts for slurry process with low acidity and good activity for hydrogenation and hydrogenolysis reactions.

[0041] As will be described further below, fresh catalyst or catalyst precursor may be fed together with bio-oil feedstock 102 into a continuous slurry reactor unit 103 as depicted in Figure 1 where they may be used in the process for catalytic deoxygenation of a feedstock of bio-oils derived from lignocellulosic biomass. The reactor unit 103 may comprise one or more reactors in series, the reactors being but not limited to continuous stirred tank reactors, bubble columns or loop reactors. In one embodiment hydrogen gas 101 may be fed to the first reactor. In otherembodiments hydrogen gas may also be fed to other reactors than the first one. In one embodiment fresh hydrothermally precipitated unsupported Mo-based catalyst or promoted Mo-based catalyst as fine solid particles may be added to the bio-oil feed tank 102 or to a separate catalyst feed tank to be combined with the bio-oil in the reactor unit feed. The promoter metal in the promoted Mobased catalyst may be Ni or Co. In another embodiment, a precursor of the Mobased catalyst such as but not limited to Mo-alcoholate may be added to the bio-oil feed tank 102 or to a separate catalyst feed tank to be combined with the bio-oil in the reactor unit feed. A sulfiding agent, such as H2S, may optionally be fed from the feed tank 102 to the reactor 103. In the reactor 103, very small Mo-sulfide particles may be expected to form, as Mo precursor may decompose and Mo may react with present H2S and form sulfide M0S2 which may be expected to be the active form catalyst in the process.

[0042] From reactor unit 103, the liquid-solid slurry product and gaseous products may be led to a separation section where heavy organic product 106, light organic product 107, aqueous liquid phase, and gas phase 108 may be separated from each other. The spent catalyst predominantly ending up in the heavy organic fraction may be circulated back to the reactor unit 103. In one embodiment, some spent catalyst may be withdrawn with the heavy fraction and corresponding amount of fresh catalyst or catalyst precursor may be added to the bio-oil feedstock. In one embodiment, heavy organic fraction may be cooled down and separated first from other products in a flash drum 104 (separation step 1) operating in the pressure domain of the reactor and in a temperature range of 150 - 250 °C. Gaseous stream from the first flash drum 104 may be led to a condenser and further to a second low pressure flash drum 105 (separation step 2) where light organic fraction 107, aqueous liquid fraction, and gas product 108 may be separated from each other. The second flash drum 105 operates in a temperature range of 20 - 50 °C and at 5 to 20 bar pressure. After the separation sequences, in one embodiment, catalyst particles may be separated from the heavy organic fraction not circulated back to the reactor unit, and the heavy and light fractions may be combined to a one organic liquid product with a dry oxygen content of less than 25 wt%, preferably in the range of 10-25 wt%.

[0043] Organic fractions from the process may be led to further upgrading e.g. by fixed bed hydroprocessing process to reduce the oxygen content of the organic fraction below 5 wt%, preferably below 1 wt% to be applied after fractionation and other conditioning as green transportation fuels (jet fuels, road gasoline, marineand road diesel fuel) or as green naphtha to be used as steam cracker feedstock. In an embodiment, the organic phase and / or the deoxygenated bio-oils contained in the light organic phase may be subjected to further hydrotreatment by hydrodeoxygenation, to produce fuels and hydrocarbons. For example, a temperature of 270 to 410 °C, and a pressure of 30 to 200 bar, and a hydrodeoxygenation catalyst may be used.

[0044] Bio-oil feedstock

[0045] An example of the characteristics ofthe bio-oil feedstockto be subjected to catalytic deoxygenation in accordance with an embodiment are presented in Table 1. The characterization results of the bio-oil, presented in Table 1, indicated the oil to be typical fast pyrolysis bio-oil from lignocellulosic forest biomass.

[0046] Table 1

[0047] <

[0048]

[0049] Catalyst preparation and characterization

[0050] Example 1E1

[0051] Aqueous solution of Co Nitrate (39.9 g), Ammonium heptamolybdate (48.0 g) and thiourea (62.5 g) in deionized water were treated with concentrated HC1 to bring the pH to 0.5. The mixture was then loaded to a Teflon-lined autoclave equipped with a magnetically coupled mechanical agitator. The reactor which was sealed, was pressurized with N2 and heated to 200 °C. The autoclave was kept at 200 °C for 15 hours and was then cooled down, flushed with pressurized N2 several times and then opened to retrieve the catalyst particles in the form of an aqueousslurry. The precipitate was filtered and washed first three times with water, and afterwards three times with technical ethanol. Then, the catalyst was dried in a ro-tavapor at 50 mbar and 60 °C and stored in a nitrogen atmosphere. 40.0 g of catalyst were recovered in the form of black solid powder containing 19.4 wt% of Co and 31.5 wt%. of Mo and 46.5 wt% of S according to elemental analysis, resulting in a molar ratio of M0C0S4.41. Volumetric particle size distribution of catalyst from example 1E1 is shown in Figure 2. BET surface area was measured to be 35 m2 / g.

[0052] Example 1E2

[0053] Aqueous solution of Co nitrate (0.53 g], ammonium molybdate (0.64 g] and thiourea (0.83 g] in deionized water were treated with concentrated HC1 to bring the pH to 0.4. The mixture was then loaded to a Teflon-lined autoclave equipped with a magnetic stirrer. The reactor which was sealed, was pressurized with N2 and heated to 220 °C. The autoclave was kept at 200 °C for 20 hours and was then cooled down, flushed with pressurized N2 several times and then opened to retrieve the catalyst particles in the form of an aqueous slurry. The precipitate was filtered and washed first three times with water, and afterwards three times with technical ethanol. Then, the catalyst was dried in vacuum at 50 °C and stored in a nitrogen atmosphere. 0.75 g of catalyst were recovered in the form of solid powder containing 16.2 wt% of Co and 36.1 wt%. of Mo and 39.9 wt% of S according to elemental analysis, resulting in a molar ratio of M0C00.73S3.31. BET surface area was measured to be 126 m2 / g.

[0054] Comparative example CE1

[0055] Aqueous solution of ammonium molybdate (8.00 g] and thiourea (13.75 g] in deionized water were treated with concentrated HC1 to bring the pH to 0.8. The mixture was then loaded to a Teflon-lined autoclave equipped with a magnetically coupled mechanical agitator. The reactor which was sealed, was pressurized with N2 and heated to 240 °C. The autoclave was kept at 240 °C for 20 hours and was then cooled down, flushed with pressurized N2 several times and then opened to retrieve the catalyst particles in the form of an aqueous slurry. The precipitate was filtered and washed first three times with water, and afterwards three times with technical ethanol. Then, the catalyst was dried in vacuum at 50 °C and stored in a nitrogen atmosphere. 6.67 g of catalyst were recovered in the form of solid powder containing 56.0 wt% of Mo and 39.4 wt% of S according to elementalanalysis, resulting in a molar ratio of M0S2.1. BET surface area was measured to be 149 m2 / g.

[0056] Comparative example CE4

[0057] Aqueous solution of Co Nitrate (0.53 g), ammonium molybdate (0.64 g) and thiourea (0.55 g) in deionized water were treated with concentrated HC1 to bring the pH to 0. The mixture was then loaded to a Teflon-lined autoclave equipped with a magnetic stirrer. The reactor which was sealed, was pressurized with N2 and heated to 240 °C. The autoclave was kept at 200 °C for 20 hours, and was then cooled down, flushed with pressurized N2 several times and then opened to retrieve the catalyst particles in the form of an aqueous slurry. The precipitate was filtered and washed first three times with water, and afterwards three times with technical ethanol. Then, the catalyst was dried in vacuum at 50 °C and stored in a nitrogen atmosphere.0.60 g of catalyst was recovered in the form of solid powder containing 9.8 wt% of Co and 45.6 wt% of Mo and 20.2 wt% of S according to elemental analysis, resulting in molar ratio of M0C00.34S1.32.

[0058] 1CP analysis

[0059] In the above examples 1 E 1, 1E2, CE1, concentrations of catalyst elements Co, Mo and S were measured with 1CP-OES method. The solids were first subjected to a microwave-assisted digestion, treating ca 0.1 g sample aliquots into the mi-crowith 9 ml HNO3 (65%) and 1 ml H2O2 (30%) in sealed tubes. Following digestion method was applied: 1) temperature was increased to 210 °C in 20 minutes, 2) temperature was held at 210 °C for 15 minutes. After the samples had been cooled down, they were diluted to 50 ml with ultrapure type 1 water. Prior to the analysis the solutions were filtered with syringe filters (0.45 pm). The elements were analysed with 1CP-OES equipment from 1 / 100 and undiluted samples. Dilutions were made with 1% nitric acid solution. Multielemental standard solutions were used as calibration standards and control samples in the 1CP-0ES measurement.

[0060] Physisorption analysis

[0061] N2 adsorption-desorption was performed with Micromeritics 3Flex 3500 tool in liquid nitrogen. Before the analysis, samples were degassed in Micromeritics VacPrep tool for a minimum of 18 hours at 120 °C.

[0062] In the above example 1E1, particle size distribution (PSD) of the catalyst was obtained with laser diffraction by using Malvern Mastersizer 3000 with Hydro LV wet sample dispersion unit utilizing water as the dispersant. The catalysts weredissolved into deionized water before inserting into the sample unit. The material properties were set assume the sample as spherical particles, and refractive index was set to 4.8323, absorption index to 0.01, and density to 1 g / cm3. A total of 15 measurements were conducted for every sample with measurement time of 10 seconds. Sonication was applied to the slurries prior to the measurement, using 40 W in-line sonication probe of the sample unit at 20% power.

[0063] According to the analytical results, catalysts of examples 1E1 and 1E2 represent typical unsupported mixed sulfide catalysts comprising both Co and Mo sulfide phases, all obtained as finely divided particles. Catalyst of comparative example CE1 did not contain Co promoter. Catalyst of comparative example CE4 contained low sulfur loading.

[0064] Catalyst tests, batch reactor

[0065] Catalyst performance and effect of catalyst composition or process temperature on observed catalyst activity and therefore process performance is evaluated below using measured hydrogen consumption in the process, measured model compound conversion to products, and measured degree of deoxygenation, as performance indicators. These are existing methods for process performance evaluation.

[0066] Example 1E3 - Catalyst test, batch reactor

[0067] Catalyst of example 1E1 (0.50 g), bio-oil (Table 1) (23.7 g), sulfiding agent DMDS (0.03 g) and dodecane (24.9 g) were loaded to the reactor (Buchi no-voclave, wetted parts Hastelloy, 200 ml). Reactor was flushed with nitrogen and pressure-tested. After successful pressure test, reactor was flushed with hydrogen and loaded to the reaction pressure of 60 bar at room temperature. Temperature and pressure were recorded and reactor mixing and heating to reaction temperature were started. Experiment time of 180 min was considered started when desired temperature of 380 °C was reached.

[0068] After desired reaction time of 180 min, reactor was cooled down by cooling down the heating jacket for 1 h. Temperature and pressure of the reactor were recorded and gas sample was taken to gas bag from headspace of the reactor. Reactor was flushed with nitrogen and opened. Liquids and solids from the reactor were collected to a sample bottle, and reactor vessel and reactor internals (stirrer, thermoelement, dip tube) were flushed with acetone overnight. The clean-up acetone was collected for mass-balance closure.The collected reactor sample was vacuum filtered and separated to light organic phase and aqueous phase in separation funnel. The formed filter cake was washed with reactor clean-up acetone and make-up acetone, dried and collected. The clean-up acetone of the reactor and filtrate from solids washing were combined, and acetone was evaporated in rotavapor at 300 mbar at 40 °C. Evaporation residue was collected as oil residue and analysed separately. H2 consumption and degree of deoxygenation are reported in Table 2a.

[0069] Comparative example CE2 - Catalyst test, batch reactor

[0070] Catalyst of example 1E1 (0.49 g), bio-oil (Table 1) (24.4 g), sulfiding agent DMDS (0.02 g) and dodecane (24.9 g) and dodecane (26.0 g) were loaded to the reactor (Buchi novoclave, wetted parts Hastelloy, 200 ml). Reactor was flushed with nitrogen and pressure-tested. After successful pressure test, reactor was flushed with hydrogen and loaded to the reaction pressure of 60 bar at room temperature. Temperature and pressure were recorded and reactor mixing and heating to reaction temperature were started. Experiment time of 180 min was considered started when desired temperature of 250 °C was reached.

[0071] After desired reaction time of 180 min, reactor was cooled down by cooling down the heating jacket for 1 h. Temperature and pressure of the reactor were recorded and gas sample was taken to gas bag from headspace of the reactor. Reactor was flushed with nitrogen and opened. Liquids and solids from the reactor were collected to a sample bottle, and reactor vessel and reactor internals (stirrer, thermoelement, dip tube) were flushed with acetone overnight. The clean-up acetone was collected for mass-balance closure.

[0072] The collected reactor sample was vacuum filtered and separated to light organic phase and aqueous phase in separation funnel. The formed filter cake was washed with reactor clean-up acetone and make-up acetone, dried and collected. The clean-up acetone of the reactor and filtrate from solids washing were combined, and acetone was evaporated in rotavapor at 300 mbar at 40 °C. Evaporation residue was collected as oil residue and analysed separately. H2 consumption and degree of deoxygenation are reported in Table 2a.

[0073] Example 1E4 - Catalyst test, batch reactor

[0074] Catalyst of example 1E2 (0.53 g), bio-oil (Table 1) (23.9 g), sulfiding agent DMDS (0.03 g) and dodecane (24.7 g) were loaded to the reactor (Buchi novoclave, wetted parts Hastelloy, 200 ml). Reactor was flushed with nitrogen andpressure-tested. After successful pressure test, reactor was flushed with hydrogen and loaded to the reaction pressure of 60 bar at room temperature. Temperature and pressure were recorded and reactor mixing and heating to reaction temperature were started. Experiment time of 180 min was considered started when desired temperature of 380 °C was reached.

[0075] After desired reaction time of 180 min, reactor was cooled down by cooling down the heating jacket for 1 h. Temperature and pressure of the reactor were recorded and gas sample was taken to gas bag from headspace of the reactor. Reactor was flushed with nitrogen and opened. Liquids and solids from the reactor were collected to a sample bottle, and reactor vessel and reactor internals (stirrer, thermoelement, dip tube) were flushed with acetone overnight. The clean-up acetone was collected for mass-balance closure.

[0076] The collected reactor sample was vacuum filtered and separated to light organic phase and aqueous phase in separation funnel. The formed filter cake was washed with reactor clean-up acetone and make-up acetone, dried and collected. The clean-up acetone of the reactor and filtrate from solids washing were combined, and acetone was evaporated in rotavapor at 300 mbar at 40 °C. Evaporation residue was collected as oil residue and analysed separately. H2 consumption and degree of deoxygenation are reported in Table 2a.

[0077] Comparative example CE3 - Catalyst test, batch reactor

[0078] Catalyst of example CE1 (0.55 g), bio-oil (Table 1) (24.0g), sulfiding agent DMDS (0.02 g) and dodecane (26.1 g) were loaded to the reactor (Buchi no-voclave, wetted parts Hastelloy, 200 ml). Reactor was flushed with nitrogen and pressure-tested. After successful pressure test, reactor was flushed with hydrogen and loaded to the reaction pressure of 60 bar at room temperature. Temperature and pressure were recorded and reactor mixing and heating to reaction temperature were started. Experiment time of 180 min was considered started when desired temperature of 380 °C was reached.

[0079] After a desired reaction time of 180 min, reactor was cooled down by cooling down the heating jacket for 1 h. Temperature and pressure of the reactor were recorded and gas sample was taken to gas bag from headspace of the reactor. Reactor was flushed with nitrogen and opened. Liquids and solids from the reactor were collected to a sample bottle, and reactor vessel and reactor internals (stirrer, thermoelement, dip tube) were flushed with acetone overnight. The clean-up acetone was collected for mass-balance closure.The collected reactor sample was vacuum filtered and separated to light organic phase and aqueous phase in separation funnel. The formed filter cake was washed with reactor clean-up acetone and make-up acetone, dried and collected. The clean-up acetone of the reactor and filtrate from solids washing were combined, and acetone was evaporated in rotavapor at 300 mbar at 40 °C. Evaporation residue was collected as oil residue and analysed separately. H2 consumption and degree of deoxygenation are reported in Table 2a.

[0080] Table 2a. Hydrogen consumption and degree of bio-oil deoxygenation for batch-wise hydroprocessing experiments

[0081] >

[0082]

[0083] As shown in Table 2a, hydroprocessing of bio-oil in presence of catalysts 1E1 and 1E2 (Examples 1E3 and 1E4, respectively), led to high hydrogen consumption as per analysis of the headspace of the reactor. Furthermore, analysis of the processed oil samples clearly indicated a high degree of deoxygenation (above 50%) achieved during slurry phase hydroprocessing. Notably, treatment of bio-oil at lower temperature (250 °C, comparative example CE2) resulted in substantially lower hydrogen consumption due to insufficient reactivity at lower temperatures. The productor comparative example CE2 could not be separated from the aqueous phase and solvent due to high oxygen content likely originating from low hydrogen consumption, therefore degree of deoxygenation could not be measured reliably. Comparative Mo based catalyst of example CE1 resulted in lower degree of bio-oil deoxygenation compared to Co-promoted catalysts 1E1 and 1E2 hereby indicating beneficial effect of Co-promotion the reaction outcome.Examples IE9 and CE5

[0084] Catalyst of examples 1E2 (0.10 g) and CE4, bio-oil model compound isoeugenol (Table 2b) (3.0 g), sulfiding agent DMDS (0.23 g) and dodecane (100 ml) were loaded to the reactor (Buchi novoclave, wetted parts Hastelloy, 200 ml). The reactor was flushed with nitrogen and pressure-tested. After a successful pressure test, the reactor was flushed with hydrogen and loaded to a reaction pressure of 30 bar at room temperature. Temperature and pressure were recorded, and heating to reaction temperature was started. When the reactor reached the desired temperature of 300 °C, mixing was started and reaction experiment was considered started.

[0085] After a desired reaction time of 180 min, a sample was taken via a diptube from the reactor liquid phase. A gas phase sample was taken to a gas bag and analysed with micro-GC for permanent gases. The liquid samples taken via the diptube along the test run were analysed with gas chromatography (GC) equipped with Agilent ULTRA 119091A-115 capillary column (50 m x 320 pm x 0.52 pm) and a flame-ionization detector (GC-F1D). Conversion of the model compound isoeugenol and the hydrogenation product dihydroeugenol was calculated, as shown in Equation 1 below.

[0086] > >

[0087]

[0088] where X is the conversion, cie, cdheare the concentrations of isoeugenol and dihydroeugenol, and cie0is the concentration of isoeugenol at the start of the experiment.

[0089] Table 2b

[0090]

[0091] The results presented in Table 2b show the importance of high sulfur content for high conversion of the bio-oil feedstock material.Examples 1E5-1E8, Catalyst tests, continuous slurry reactor

[0092] Products of examples IE5-IE8 were produced in a continuous slurry hydrocracker plant described previously in publication Bergvall, N., et al., Corefining of Fast Pyrolysis Bio-Oil with Vacuum Residue and Vacuum Gas Oil in a Continuous Slurry Hydrocracking Process. Energy & Fuels, 2020. 34(7): p. 8452-8465. Catalyst of example 1E1 was mixed with bio-oil described in Table 1 and was then fed to the process targeting bio-oil to catalyst ratio of 1:200, 5000 ppm by weight.

[0093] The reactor was loaded with diesel as an inert compound for the heating of the reactor before starting feeding of the bio-oil slurry. The full setup was pressure tested with nitrogen. The heating and stirring of the reactor and hydrogen feeding were started with diesel-containing reactor. The catalyst slurry of hydrocarbon solvent and catalyst in known amount, sulfiding agent DMDS and bio-oil were weighed to the slurry tank from separate storage vessels aiming at catalyst loading 1500 ppm based on molybdenum content. Once the reactor reached the desired temperature and feeding from slurry tank was started.

[0094] The reactor was assumed stable and representative after three reactor replacements (4.5 h with the residence time of 1.5 h) in sampling temperature. The product tanks were emptied, and sampling period of three reactor replacements (4.5 h with the residence time of 1.5 h) was started. Sampling was carried out by collecting the sample collected to the heavy product tank as such, containing aqueous and oil phase. The light products were collected to the separator and was sampled by emptying the separator to the product tank and measuring the weight after sampling period.

[0095] Light samples were received from LP-LT separator contained a light aqueous phase and light oil phase. These phases were separated in a separation funnel at room temperature and pressure and referred as light oil phase and light aqueous phase. Major part of the products was received as bottom product from the HP-HT separator and contained aqueous phase and oil phase, referred as heavy aqueous phase and heavy oil phase. The samples were separated by centrifuging in room temperature.

[0096] The results of the test runs are presented in Tables 3 to 5.Table 3. Conditions, mass balance and product phase distribution of the continuous test runs of examples 1E5-1E8

[0097]

[0098] Oil phase elemental characterization results are presented in Table 4. Physical properties of oil products and the properties of oil phase are presented in Table 5. One can see that all examples 1E5-1E8 resulted in a substantial deoxygenation of the bio-oil product, i.e. lower oxygen and water content. While the reactor temperature was maintained in the range of desired high values, it can be seen that higher reaction temperature was favouring deoxygenation process thereby result-ing in less oxygenated bio-oil products.

[0099] Table 4. Elemental analyses and water content in feed oil and organic products (wt%)

[0100] <

[0101] <

[0102] <

[0103] <

[0104]

[0105] < Table 5. Feed oil and product oil phase physical properties

[0106]

[0107] * CAN denotes the carboxylic acid number, i.e. the amount of carboxylic acid groups in product oil, measured as KOH consumption in titration.

[0108] The water content was analyzed by Karl Fischer titration using a Metrohm 795 KFT Titrino titrator (ASTM E 203). Elemental composition (CHN) was analyzed using an Elementar VAR10MAX CHN analyzer (ASTM D 5291). Oxygen was calculated as difference from carbon, hydrogen, nitrogen and water content analyses. The carboxylic acid number (CAN) was determined with a 785 DMP Titrino analyzer (modified ASTM D 664). Kinematic viscosity and density at 40 °C were measured with Anton Paar Stabinger Viscometer SVM 3000.

[0109] Oxygen content (0 (wt-%)) was calculated as difference from carbon, hydrogen, nitrogen and water content analyses, as described below.

[0110] (0 (wt-%) = 100%-C(wt-%)-H(wt-%)-N(wt-%)-(16 / 18)*H20(wt-%)).

[0111] Degree of deoxygenation was calculated from organic oxygen content of the feed and product oils, as described below.

[0112] Degree of deoxygenation (DoD) was calculated as

[0113] DoD = 100%-(0bio-oil feed(wt-%) / Ooil product(wt-%))

[0114] Results of the oxygen content analysis are presented above in Table 4. Results of the degree of deoxygenation analysis are presented above in Table 2a. Results of the carboxylic acid number (CAN) analysis are presented above in Tables 1 and 5.

[0115] Thus, in an embodiment, bio-oil feedstock may be upgraded into an oil product usable for the production of advanced transportation fuels and / or green chemical intermediates. For example, the bio-oil product may have a lower oxygenand water content, as well as lower kinematic viscosity, and density, compared to the bio-oil feed.

[0116] It will be obvious to a person skilled in the art that, as the technology advances, the inventive concept can be implemented in various ways. The inven-tion and its embodiments are not limited to the examples described above but may vary within the scope of the claims.

Claims

CLAIMS1. A process for catalytic deoxygenation of a feedstock of bio-oils derived from lignocellulosic biomass, comprising:hydrotreating a slurry comprising the feedstock of bio-oils derived from lignocellulosic biomass and an unsupported molybdenum-based catalyst with hydrogen ata temperature of 270 to 410 °C, preferably 350 to 400 °C, more preferably 370 to 390 °C, and at a pressure of 30 to 200 bar,wherein the unsupported molybdenum-based catalyst comprises a promoter metal (Me) and has the formula MoxMeySz, wherein x=l, y=0.3 to 1.5, z=3 to 5, and Me is Co or Ni.

2. A process according to claim 1, wherein feedstock of bio-oils derived from lignocellulosic biomass comprises bio-oils obtained by pyrolysis, preferably fast pyrolysis, or liquefaction of lignocellulosic biomass.

3. A process according to claim 1 or 2, whereinx=l, y=0.5 to 0.8, z=3 to 4.

4. A process according to any one of the claims 1 to 3, wherein the concentration of the unsupported molybdenum-based catalyst in the slurry is 100 to 10000 ppm by weight, preferably 200 to 5000 ppm by weight.

5. A process according to any one of the claims 1 to 4, wherein the solid content of the slurry is less than 5 wt-%, preferably less than 1 wt-%, more preferably less than 0.1 wt-%.

6. A process according to any one of the claims 1 to 5, wherein the unsupported molybdenum-based catalyst has a particle size D50 of 5 to 100 microns.

7. A process according to any one of the claims 1 to 6, wherein the pressure is 50 to 200 bar, preferably 60 to 200 bar.

8. A process according to any one of the claims 1 to 7, wherein the hydrotreating is performed in a slurry reactor.

9. A process according to any one of the claims 1 to 8, comprising continuously replacing unsupported molybdenum-based catalyst spent during hydrotreating the slurry with fresh amounts of said unsupported molybdenum-based catalyst.

10. A process according to any one of the claims 1 to 9, further comprising a separation step where a gas phase and an aqueous phase formed during hydrotreating the slurry are separated from an organic liquid phase of the slurry and the gas phase is optionally partly circulated back to the slurry.

11. A process according to claim 10, further comprising a step wherelight and heavy organic phases are separated from each other and the heavy organic phase containing catalyst is optionally partly or fully circulated back to the slurry.

12. A process according to claim 10, wherein the spent catalyst is re-moved from the slurry with the part of the heavy organic phase not circulated back to the slurry, and fresh catalyst is added to the slurry.

13. A composition of deoxygenated bio-oils obtainable by the process according to any one of the claims 1 to 12.

14. A composition according to claim 13, having- a kinematic viscosity of 35 mm / s2or less, preferably in the range of 4 to 25 mm / s2- a density of 1.06 g / ml or less, preferably in the range of 1 to 0.9 g / ml - a dry oxygen content of less than 25 wt%, preferably in the range of 10 to 25 wt%.

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