CATALYSTS AND SELECTIVE PROCESS FOR THE PRODUCTION OF RENEWABLE AVIATION FUELS AND BIOFUEL PRODUCED

Abstract

The present invention relates to a process for converting vegetable oils, animal fats, residual edible oils and carboxylic acids into renewable liquid fuels, such as bionaphtha, bioJET-A1 and renewable diesel, for use in a mixture with fossil fuels. The process consists of two steps: hydrotreating and hydrocracking. The effluent from the hydrotreatment step presents aromatics, olefins and compounds resulting from the polymerization of esters and acids in its composition. This fact occurs due to the use of partially reduced catalysts and without injection of sulfide agent and allows obtaining a bioJET-A1 with adequate quality for use in a mixture with fossil kerosene. At the same time, the process generates, in addition to products in the distillation range of naphtha, kerosene and diesel, high molecular weight linear paraffins (with up to 40 carbon atoms).

Claims

1. A selection process for the production of renewable aviation fuels characterized in that: a) it is carried out in two stages; b) the renewable feedstock of the first stage is selected among vegetable oils, animal fats, residual edible oils or acids; c) the first stage (as described in BR 102019027610-0) is active for hydrodeoxygenation and polymerization reactions, with the formation of hydrocarbons with longer carbon chains than those present in the feedstock; d) the effluent from the first stage, containing n-paraffins, olefins and aromatics and free of organosulfurates, is directed to the second stage of reaction, without the need for purification to eliminate contaminants; e) the second stage of the conversion presents multiple catalytic beds, for hydroisomerization and hydrocracking reactions, whose catalysts have different chemical and physical characteristics; f) the catalytic beds of the second stage show gradation of activity; g) the operating conditions of the second stage are: pressure in the range of 3 MPa to 8 MPa, average temperature of the catalytic bed between 240° C. and 380° C. and hydrogen/feedstock ratio between 100 and 600 NL H.sub.2/L of feedstock; h) the operating conditions of the second stage may be the same or different from those of the first stage.

2. The process according to claim 1, characterized in that it comprises the following steps: a. the composite feedstock stream (1) is mixed with a recycled product stream from the separator vessel (V1) and receives an injection of recycle gas, rich in hydrogen; a fraction of the composite feedstock and recycle gas is heated and directed to the inlet of the first stage reactor (R1); b. the fraction of unheated composite feedstock and recycle gas is directed to the region between the catalytic beds of the reactor (R1), to control the reactor temperature; c. the effluent from the reactor (R1) is directed to the separator vessel (V1) where an aqueous phase, a gaseous phase, composed of light hydrocarbons, and a liquid phase are separated; d. a fraction of the liquid phase obtained in (d) is returned to the beginning of the process to dilute the composite feedstock; e. another fraction of the liquid phase obtained in (d) is mixed with the replenishment hydrogen (4) and directed to the second stage reactor (R2), located downstream the reactor (R1); f. the effluent from the reactor (R2) is sent to the separator vessel (V2), from where the gaseous stream rich in H.sub.2 is recycled directly to the reactor (R1), without the need for treatment in amine units to remove contaminants; g. the liquid stream from the separator vessel (V2) is directed to the distillation tower (T1) where a lighter stream (6) is separated, rich in olefins with 3 to 4 carbon atoms; a stream of gasoline (7), rich in isomers and olefins; a stream specified as bioJET-A1, according to ASTM 7655 standard and a renewable diesel stream (9).

3. The process according to claim 2, characterized in that the vegetable oil is selected from the group consisting of: castor, soy, canola, peanut, palm (dendê) and babassu oil, pure or mixed in any ratio.

4. The process according to claim 2, characterized in that the renewable feedstock is animal fat of any origin.

5. The process according to claim 4, characterized in that the feedstock is a mixture of vegetable oil and animal fat in any ratio.

6. The process according to claim 2, characterized in that the first catalytic bed of the second stage reactor (R2) consists of a catalyst with high activity for hydroisomerization reactions and low activity for hydrocracking reactions and subsequent beds consisting of catalysts of hydrocracking, with a progressive increase in activity.

7. Catalysts for the production of renewable aviation fuels, used in the second stage of the process described in claim 1, characterized in that: a. they are composed of group VIIIB metal oxides (Type 3 catalyst), mainly platinum, in concentrations of 0.1 to 1.0% in mass, preferably between 0.3 and 0.5% in mass, completely reduced in the process conditions and not presenting compounds with different oxidation states, as occurs with the first stage catalysts; b. the first catalyst bed has high activity for hydroisomerization reactions and low activity for hydrocracking reactions; c. the other beds show a progressive increase in hydrocracking activity; d. the gradation of hydrocracking activity is obtained through the selection of supports containing molecular sieves; e. the molecular sieve content in the catalyst particle is in the range of 20% to 70% in mass.

8. The catalysts for the second stage according to claim 7, characterized in that the catalyst support of the first bed is a molecular sieve selected from the group formed by SAPO-11, SAPO-31 and SAPO-41.

9. The catalysts for the second stage according to claim 7, characterized in that the support of the subsequent beds are zeolite-type molecular sieves, selected from the group formed by Beta zeolite, Y zeolite, ZSM-22, ZSM-23 and ZSM-35.

10. The catalysts for the second stage according to claim 7, characterized in that the support of the subsequent beds are Beta zeolite, ZSM-22, ZSM-23 and ZSM-35.

11. The catalysts for the second stage according to claim 7, characterized in that the molecular sieve content in the catalyst particle is 30% to 50% in mass.

12. The catalysts for the second stage according to claim 7, characterized in that they are prepared in the form of particles with an equivalent diameter of 1 mm to 5 mm.

13. A biofuel produced according to the process of claim 1 and using the catalysts described in claim 7, characterized in that it is a stream specified as bioJET-A1, in accordance with ASTM 7655 standard.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0034] FIG. 1 presents a possible basic scheme for the process of the present invention without, however, limiting it to this configuration. The proposed catalytic arrangement can be summarized in the use of two stages.

[0035] In FIG. 1, the process feedstock (1) is pumped by pump B1 and mixed with the recycle from pump B2. The feedstock is mixed with hydrogen from the recycle compressor K1. Part of the feedstock can be injected between the catalytic beds of the reactor R1, and the rest is heated in the heat exchanger batteries and in the furnace F1 and fed into the reactor inlet. Reactor R1 consists of several catalyst beds containing Type 1 and Type 2 catalysts, as described in BR 102019027610-0. The product from reactor R1 is sent to separator vessel V1 where the different phases are separated. The aqueous phase (5), rich in CO.sub.2, can be directed to the refinery process water; the gaseous phase, consisting of light hydrocarbons, with the average boiling point of gasoline and unreacted hydrogen, is cooled and mixed with the cooled effluent from reactor R2. Part of the liquid hydrocarbon stream from V1 is mixed with replenishment hydrogen (4) and fed into reactor R2. Most of the stream is pumped by pump B2 to make up the dilution stream of the unit feedstock stream. The reactor R2, consisting of one or more beds of Type 3 catalysts, object of the present invention, is responsible for the generation of lower boiling point products. The effluent from reactor R2 is cooled, mixed with the gaseous stream separated from V1 and sent to phase separator V2. In separator V2, the gaseous stream rich in H.sub.2 is recycled to reactor R1 and the hydrocarbon stream (3) is sent to distillation tower T1, where there are separated: a light stream (6), rich in C3 and C4 olefins; a stream in the gasoline distillation range (7), rich in isomers and olefins; a stream (8) in the bioJET-A1 distillation range and that meets the specifications of ASTM 7655 standard; and the stream (9) in the diesel distillation range. The stream (10), heavier than diesel, can be recycled to reactor R2 to be converted or processed into HCC units.

DESCRIPTION OF THE INVENTION

[0036] The process presents a solution to increase the yield of renewable kerosene (bioJET-A1) from paraffinic streams derived from biomass. The current process describes the conversion of the product generated by processing vegetable oils, animal fats, residual edible oils and carboxylic acids, as described in BR 102019027610-0, into renewable kerosene that meets the specifications of ASTM 7566 standard. In a broad way, the process of this invention proposes the conversion of biomass in two stages.

[0037] In the first stage of the process, a feedstock consisting of triacylglycerides, in the ratio between 1 and 75% by mass, and hydrocarbons, from the recycling of part of the generated product, in the ratio of 99 to 25% in mass, is hydrotreated. It should be noted that the cargo cannot contain any sulfur compounds (H.sub.2S, DMDS, etc.). The composite feedstock, after injecting a stream of hydrogen, is directed to the reactor where the hydrodeoxygenation reactions take place, in the presence of a partially reduced catalyst and without the addition of sulfide compounds. The hydrotreatment conditions are: operating pressure from 4 MPa to 10 MPa, average temperature of the catalytic bed between 320° C. and 400° C., space velocity from 0.5 h.sup.-1 to 2 h.sup.-1 and hydrogen: feedstock ratio ranging from 200 NL of hydrogen/liter of feedstock to 1000 NL of hydrogen/liter of feedstock.

[0038] In the first stage, one or more reactors containing at least one catalyst bed of group VIB metal oxide catalyst, partially reduced, supported on materials with high specific area and low cracking activity (Type 1 catalyst, as described in document BR 102019027610-0). This catalytic system favors hydrodeoxygenation and polymerization/oligomerization reactions. The Type 1 catalyst beds may or may not be followed by other catalyst beds containing metal oxide catalysts from group VIB and VIIIB (e.g.: NiMo, NiW), partially reduced, supported on materials with high specific area and acidity, in order to promote cracking reactions (Type 2 catalyst, as described in document BR 102019027610-0). The operating conditions, mainly the reaction temperature, are adjusted so that the products formed in this first stage have molecules with a carbon number of 19, which corresponds to the final boiling point (ASTM 2887 analysis) of approximately 320° C. Therefore, to obtain products in the distillation range characteristic of bioJET-A1 (ASTM 7566) and renewable diesel, the final stream of the process needs to be distilled.

[0039] The produced renewable hydrocarbon stream has a different composition from those produced in similar processes because it contains olefinic and aromatic compounds, in addition to normal paraffins. In the bioJET-A1 stream, the high concentration of olefins leads to low oxidation stability, and the high concentration of normal paraffins implies a high freezing point, making its use unfeasible as a fuel according to the specifications of the ASTM 7566 standard. Thus, the distillate cut needs to be hydrogenated to reduce the concentration of olefins and aromatics; however, to frame the freezing point at values below -40° C., the stream needs to be isomerized to generate branched aliphatic compounds. These reactions occur in the second stage of the process, which is the object of the present invention.

[0040] Due to the absence of contaminants, such as organosulfur compounds, the bioJET-A1 cut generated in the first stage can be sent to the second stage without the need for purification processes.

[0041] The second stage of conversion, responsible for increasing the yield of bioJET-A1, is characterized by using multiple catalyst beds, preferably from three to five, whose catalysts have different chemical and physical characteristics.

[0042] The second stage catalyst beds are composed of group VIIIB metal oxide catalysts (Type 3 catalyst), mainly platinum, in concentrations of 0.1 to 1.0% in mass, preferably between 0.3 and 0.5% in mass, completely reduced under process conditions and not showing compounds with different oxidation states as occurs with first stage catalysts (Type 1 and Type 2 catalysts).

[0043] The second stage catalyst beds should show a gradation in hydrocracking activity. The first catalytic bed has high activity for hydroisomerization reactions and low activity for hydrocracking reactions. The other beds should show a progressive increase in hydrocracking activity. The gradation of activity is obtained through the selection of supports containing molecular sieves, such as SAPO-11, SAPO-31 and SAPO-41, recommended for the initial bed, or as zeolites: Beta, ZSM-22, ZSM-23 and ZSM-35 for subsequent beds. These catalysts should be prepared in the form of particles with an equivalent diameter of 1 mm to 5 mm.

[0044] The molecular sieve content in the catalyst particle may vary from 20% to 70% in mass, preferably from 30% to 50% in mass, for a more precise adjustment of the catalyst activity.

[0045] The activity of molecular sieves is a function of the strength of the acidic active sites and can be controlled by varying the Si/Al ratio (SAR) present on the catalyst support. Different Si/Al ratios can be adjusted depending on the binder used, which may be gamma-alumina, silica, clays or kaolin, but not limited to these compounds.

[0046] The different catalysts of the second stage of conversion can operate under the same operating conditions or under different conditions, mainly with reduced pressure and/or increased temperature, in order to favor hydrocracking reactions. The operating conditions are: pressure from 3 MPa to 8 MPa; average temperature of the catalytic bed between 240° C. and 380° C. and hydrogen/feedstock ratio between 100 and 600 NL H.sub.2/L of feedstock.

[0047] Based on the knowledge presented, it is ascertained that a possible sequence for the arrangement of the second stage catalytic beds, without restriction to other arrangements, aiming at increasing the selectivity for the production of bioJET-A1, from a stream containing n-paraffins, olefins and aromatics (according to the effluent stream of the first stage of conversion) is: a first catalytic bed with the function of promoting mono branching and with low cracking activity, composed of Pt supported on SAPO-11, ZSM-22 with low Si/Al ratio or Beta zeolite passivated with organic acids; a second catalytic bed aiming at increasing the number of branches, composed of Pt supported on mesoporous sieves (e.g.: Pt/MCM-41); a third catalytic bed designed to promote small cracking and increased isomerization, composed of Pt catalyst with low acidity (e.g., Pt/ZSM-22) and a fourth catalytic bed to promote moderate cracking with some isomerization, composed of Pt catalyst Pt with medium acidity (e.g., Pt/Beta Zeolite).

[0048] For a better use of the characteristics of the various formulations of catalysts, a two-stage process scheme is proposed, without, however, limiting the same to this configuration.

[0049] In the first stage of the conversion, a Type 1 catalyst is used (as described in BR 102019027610-0), active for hydrodeoxygenation and polymerization reactions, with the formation of hydrocarbons with carbon chains longer than those present in the feedstock, composed of group VIB metal oxides, mainly Mo and W, supported on materials with a high specific area and high porosity, one of the most used materials being y-alumina (γ-Al.sub.2O.sub.3) with a specific area between 200 and 400 m.sup.2/g and pore volume from 0.5 to 1.0 cm.sup.3/g. In addition to providing a high specific area, in which the active components are dispersed in the form of small particles, the support provides mechanical strength and thermal stability, preventing the sintering of the catalyst inside the reactor.

[0050] Still in the first stage of the conversion, it is optional to use a Type 2 catalyst (as described in BR 102019027610-0), composed of metal oxides of groups VIB and VIIIB (preferably Ni), which are usually bimetallic in the form of metal oxides (Ni—Mo, Co—Mo, Co—W and Ni—W) deposited on a support with acidic properties and active for hydrocracking reactions, such as zeolite-type molecular sieves, preferably Beta zeolite, ZSM-22, Y zeolite, etc. The function of the Type 2 catalyst is to increase the efficiency of the stream with the boiling point of the bioJET-A1 (Final Boiling Point (FBP) = 300° C.).

[0051] The Type 1 catalyst promotes the removal of oxygen atoms, preserves the unsaturation present in the feedstock and catalyzes polymerization reactions, which results in products with molecules containing a greater number of carbon atoms than the carboxylic acids present in the feedstock.

[0052] Between the first and second stages, there may be, optionally, a separation of gaseous products, with the aim of recovering the light olefinic products (of greater commercial value) and the generated water.

[0053] The second stage of the process uses at least one bed of Type 3 catalyst to promote the other hydroconversion reactions. This is the fundamental step of the invention, where it is demonstrated that greater yields of renewable kerosene are obtained through an arrangement containing multiple catalytic beds, preferably from 3 to 5 beds, with different chemical and physical characteristics, so that the first catalytic bed should present a high activity for hydroisomerization reactions and low activity for hydrocracking reactions. The following catalyst beds should show a progressive increase in hydrocracking activity.

[0054] FIG. 1 presents a basic scheme of the process for better use of the characteristics of the various formulations of catalysts, without, however, limiting the same to this configuration.

[0055] In FIG. 1, the process feedstock (1), composed of carboxylic acids, esters (triacylglycerides) of vegetable oils and animal fats, etc. is pumped by pump B1 and mixed with the recycle from pump B2. The recycle stream aims at reducing the temperature of reactor R1 due to the exothermicity of the reaction. The dilution volumetric ratio is 2 to 10 (diluent/feedstock), preferably 3 to 6. The composite feedstock is mixed with hydrogen from the recycle compressor K1. The ratio between the hydrogen flow and the combined feedstock is 200 to 800 Nm.sup.3/m.sup.3, preferably 300 to 500 Nm.sup.3/m.sup.3. Part of the feedstock can be injected between the catalytic beds of the reactor R1, and the rest is heated in the heat exchanger batteries and in the furnace F1 and fed into the reactor inlet. The reactor R1 is divided into several beds containing Type 1 and Type 2 catalysts and, in order to control the reactor outlet temperature, hydrogen streams and/or a liquid stream are injected between beds. The average temperature of reactor R1 ranges from 300 to 400° C., preferably from 320 to 360° C. The product from reactor R1 is sent to separator vessel V1, where the different phases are separated. The aqueous phase (5), rich in CO.sub.2, can be directed to the refinery process water; the gaseous phase, composed of light hydrocarbons, with the average boiling point of gasoline and unreacted hydrogen, is cooled and mixed with the cooled effluent from reactor R2. The liquid hydrocarbon stream from V1 is mixed with replenishment hydrogen (4) and fed into reactor R2, with most of the stream being pumped by pump B2 to make up the unit feedstock dilution stream. The reactor R2, consisting of one or more beds of Type 3 catalysts, is responsible for generating lower boiling point products. The effluent from reactor R2 is cooled, mixed with the gaseous stream separated from V1 and sent to phase separator V2. In separator V2, the gaseous stream rich in H.sub.2 is recycled to reactor R1 and the hydrocarbon stream (3) is sent to the distillation tower T1, where the lighter stream (6), rich in C3 and C4 olefins, can be sent to the gas recovery unit of the FCC unit, or sold directly. The stream (7) is composed of gasoline rich in isomers and olefins. The stream (8) is bioJET-A1 distillation range and meets the specifications of the ASTM 7655 standard. Stream (9) consists of renewable diesel that can be added directly to the diesel pool.

EXAMPLES

[0056] In the tests carried out to determine the yields of bioJET-A1, a reactor with a volume of 5 cm.sup.3 was used, containing Type 1 catalyst (transition metal oxide, partially reduced and supported on porous solid of low acidity), operating isothermally. As feedstock, vegetable oil (tests 1 and 2) and animal fat (tests 3 and 4) were used. After processing, the gaseous stream, containing CH.sub.4, CO.sub.2 and C.sub.3H.sub.8, and the water were separated and the stream composed of saturated and olefinic hydrocarbons was used to prove the proposed innovation in the production of bioJET-A1. The liquid product was cooled and analyzed by gas chromatography coupled with a mass spectrometry detector (GC-MS) to identify the compounds. A capillary column model AC210173.038 measuring 40.0 m x 100 .Math.m x 0.20 .Math.m was used. The distillation curve of the product was determined from the boiling point of normal paraffins in a similar way to the methodology used in the ASTM 2887 standard for determining the simulated distillation curve. The points of the ASTM D86 distillation curve, as required by the ASTM 7566 standard, were estimated from the simulated distillation curve obtained and converted through the correlations presented in “Analytical Correlations Interconvert Distillation Curve Types”, Oil&Gas Journal, vol 84, 1986, August 25, pp 50-57. The typical composition of the generated liquid streams, for soy oil and animal fat feedstocks, are shown in Table 1.

TABLE-US-00001 Test 1 Test 2 Test 3 Test 4 Feedstock Soybean oil Soybean oil Fat Fat Pressure MPa 60 60 60 40 Temp °C 350 355 330 330 H2feedstock Nm.sup.3/m.sup.3 607 600 562 601 WHSV ASTM 2887 - (%mass) h.sup.-1 1.50 0.80 1.50 1.50 IBP 193.2 190.6 195.4 196.4 10 291.4 284.6 283.4 282.8 30 313.6 306.6 294.2 298.2 50 346.8 315.8 315.0 314.4 70 485.4 330.0 319.2 318.0 90 550.2 486.6 462.2 488.0 FBP 610.2 601.2 580.4 598.4 % JET-A1 150-300 14 20 31 29 % Diesel 300-350 34 47 50 44 % Light oil 350-450 13 6 3 5 % Heavy oil 450+ 33.0 17.3 10.3 16.0 % Olefins JET-A1 81.8 68.2 34.4 37.2 % n-paraffin JET-A1 18.2 31.8 65.6 62.8 % Olefins diesel 63.6 51.1 22.1 33.1 % n-paraffin diesel 36.4 48.9 77.9 66.9

[0057] To evaluate the catalytic activity of the second stage of conversion, the product of test 2 in Table 1 was used. The chromatographic analysis of this stream (without the heavy oil, which was not analyzed due to the limitations of the employed chromatographic method) is presented in Table 2, highlighting the relevant amounts of olefinic and aromatic compounds, characteristic of the first stage of conversion.

TABLE-US-00002 No. of carbons n-Paraffins (% mass) iso-Paraffins (% mass) Olefins (% mass) Alkylaromatics (% mass) C11 0.0 0.0 0.0 0.0 C12 0.0 0.0 0.0 0.0 C13 0.1 0.0 0.1 0.0 C14 0.2 0.1 0.2 0.1 C15 0.9 0.2 0.4 0.1 C16 8.9 0.6 2.6 0.1 C17 7.6 1.3 1.8 0.0 C18 33.7 6.5 20.4 1.2 C19 0.5 0.3 7.7 0.7 C20 0.8 0.3 1.5 0.1 C21 0.2 0.2 0.6 0.0 C22 0.1 0.0 0.1 0.0 C23 0.0 0.0 0.0 0.0 C24 0.0 0.0 0.0 0.0

[0058] The liquid effluent from the tests presented in Examples 1 to 3, to evaluate the catalytic activity of the second stage of conversion, was cooled and analyzed by gas chromatography coupled with a mass spectrometry detector (GC-MS) to identify the composition. A capillary column model AC210173.038 measuring 40.0 m x 100 .Math.m x 0.20 .Math.m was used. The distillation curve of the product was determined from the boiling point of n-paraffins, similarly to the methodology used in the ASTM 2887 standard for the determination of the simulated distillation curve. The freezing point was calculated based on thermodynamic calculations adjusted to experimental data, as taught by Reddy, S. R.; Fuel, 1986, December, 1647-1652 – “A thermodynamic model for predicting n-paraffin crystallization in diesel fuels”. The points of the ASTM D86 distillation curve, as determined by the ASTM 7566 standard, were estimated from the simulated distillation curve obtained and converted through the correlations presented in “Analytical Correlations Interconvert Distillation Curve Types”, Oil&Gas Journal, vol 84, 1986, August 25, pp 50-57. The distillation curve of the product was obtained from the additivity of the distillation point of the pure compounds weighted by the inverse of their respective mass fraction.

Example 1

[0059] Test with CAT1 (Type 3): the support was prepared with 50 %m of aluminosilicophosphate SAPO-11 and 50%m of gamma alumina, in cylindrical particles with about 1.3 mm in diameter. The support was dried at 120° C. for approximately 16 hours and calcined for 3 hours at a temperature of 550° C. The calcined particles were impregnated with a solution of tetraaminplatin chloride (Pt(NH.sub.3).sub.4Cl.sub.2), using the wet spot technique (0.5% in mass of Pt), and subsequently dried at 110° C. for 16 hours and calcined at 200° C. for 1 hour. The SAPO-11 molecular sieve was produced using the methodology compatible with those found in the literature, and its characterization through physicochemical methods presented structural and textural properties in accordance with those expected according to the reference http://izasc-mirror.la.asu.edu/fmi/xsl/IZA-SC/ft.xsl.

[0060] The reactor was loaded with the CAT1 catalyst (Type 3) and had the temperature raised to 400° C., with hydrogen flow, to reduce platinum (Pt). After reducing the temperature to the test condition, the feedstock of n-paraffins (from test 2 in Table 1) was injected. The tests were performed with temperature variation and WHSV = 1.4 h.sup.-1; H2/feedstock = 615 NL/L and pressure = 6 MPa. The JET-A1 yield was determined by the final boiling point of 300° C. and the initial boiling point compatible with the calculated flash point. The test results are shown in Table 3.

TABLE-US-00003 Product yields ASTM Specification 7566 a3 Test 1 Test 2 Test 3 Temperature °C 320 340 360 % naphtha % vol 0.12 1.9 4.5 % JET-A1 % vol 4.0 9.6 37.5 % diesel % vol 96.3 89.2 59.7 JET-A1 cut properties Freezing point °C -40 -35.1 -46.3 -67 Distillation (ASTM D86) T10 °C 205 max. 165.2 148.8 140.6 T50 °C 234.9 234.1 227.8 T90 °C 281.7 272.6 280.2 FBP °C 300 max. 300.0 300.0 300.0 T90-T10 °C 22 min. 115.3 120.2 145.4 Density @ 15° C. 0.730/0.770 0.7526 0.7530 0.7507 Flash point °C 38 52.4 40.8 38.0

[0061] Upon analyzing the results of Table 3, it appears that CAT1 (Type 3) has a low yield in the kerosene distillation range, although this cut presents a freezing point that is much lower than that specified by the ASTM standard. However, for tests with low conversions (Test 1), the freezing point is higher than specified.

Example 2

[0062] Test with CAT2 (Type 3): the support was prepared with 30%m of Beta zeolite and 70%m of gamma alumina, in cylindrical particles with about 1.3 mm in diameter. The support was dried at 120° C. for approximately 16 hours and calcined for 3 hours at a temperature of 550° C. The calcined particles were impregnated with a solution of tetraaminplatin chloride (Pt (NH.sub.3) .sub.4Cl.sub.2), using the wet spot technique (0.5% in mass of Pt), and subsequently dried at 110° C. for 16 hours and calcined at 200° C. for 1 hour. Beta zeolite was produced using methodology compatible with those found in the literature and its characterization, through physical-chemical methods, presented results compatible with those expected according to the reference http://izasc-mirror.la.asu.edu/fmi/xsl/IZA-SC/ft.xsl.

[0063] The reactor was loaded with a bed of CAT2 catalyst (Type 3) and had the temperature elevated, with hydrogen flow, up to 400° C. for the reduction of platinum (Pt). After reducing the temperature to the test condition, the feedstock of n-paraffins (from test 2 in Table 1) was injected. The tests were performed with temperature variation and keeping the following parameters fixed: WHSV = 1.9 h.sup.-1; H2/feedstock = 710 NL/L and pressure = 7.5 MPa. The JET-A1 yield was determined by the final boiling point of 300° C. and the initial boiling point compatible with the calculated flash point. The test results are shown in Table 4.

TABLE-US-00004 Product yields Units ASTM Specification 7566 a3 Test 1 Test 2 Test 3 Temperature °C 300 310 320 % naphtha % vol 6.8 13.5 42.6 % JET-A1 % vol 16.4 22.4 65.8 % diesel % vol 78.5 67.1 0.0 JET-A1 cut properties Freezing point °C -40 -48.3 -49.3 -38.5 Distillation (ASTM D86) T10 °C 205 max. 147.0 148.9 148.3 T50 °C 229.7 226.6 208.6 T90 °C 281.0 278.1 282.5 FBP °C 300 max. 300.0 300.0 300.0 T90-T10 °C 22 min. 138.2 134.5 157.7 Density @ 15° C. 0.730/0.770 0.7500 0.7529 0.7521 Flash point °C 38 39.5 40.9 40.5

[0064] Upon analyzing the results in Table 4, it appears that CAT2 has a high yield in the kerosene distillation range with a freezing point lower than that specified by ASTM 7566 standard. However, it is observed that there was an excessive increase in naphtha formation and decrease in the yield of diesel oil when compared to the yields obtained with CAT1 (Example 1).

Example 3

[0065] The reactor was loaded with two catalytic beds: the first containing 75% in mass of CAT1 and the second containing 25% in mass of CAT2 and had its temperature raised to 400° C. for reduction of platinum (Pt). After reducing the temperature to the test condition, the feedstock of n-paraffins (from test 2 in Table 1) was injected. The tests were performed with temperature variation and keeping the following parameters fixed: WHSV = 1.8 h.sup.-1; H.sub.2/feedstock = 700 NL/L and pressure = 6.0 MPa. The JET-A1 yield was determined by the final boiling point of 300° C. and the initial boiling point compatible with the calculated flash point. The test results are shown in Table 5.

TABLE-US-00005 Units ASTM Specification 7566 a3 Test 1 Test 2 Temperature °C 340 360 % naphtha % vol 1.4 4.3 % JET-A1 % vol 10.5 51.1 % diesel % vol 88.7 46.8 JET-A1 cut properties Freezing point °C -40 -46.4 -54.2 Distillation (ASTM D86) T10 °C 205 max. 148.7 152.1 T50 °C 231.9 217.3 T90 °C 279.7 283.8 FBP °C 300 max. 300.0 300.0 T90-T10 °C 22 min. 132.4 147.7 Density @ 15° C. 0.730/0.770 0.7521 0.7538 Flash point °C 38 40.8 43.1

[0066] Upon analyzing the results in Table 5, it is ascertained that the combination of CAT1 and CAT2 results in high product yield in the kerosene distillation range, with a freezing point lower than that specified by ASTM 7566 and with low naphtha yield. The kerosene yield obtained in test 2 of Table 5 exceeds the values obtained with the catalysts separately, as can be seen by comparing the best kerosene yields to the same (approximate) naphtha yield shown in Table 3 (Test 3) and in Table 4 (Test 1). These results demonstrate the advantage of using the combination of catalysts proposed in this invention to increase the yield of renewable kerosene, with cold flow properties (freezing point) suitable for use as aviation fuel.