PROCESS FOR OBTAINING AROMATICS AND AROMATIC STREAM

Abstract

The present invention addresses to a process for the production of aromatic compounds from streams containing linear chains with 5 to 18 carbon atoms, of fossil or renewable origin, and application in the field of catalytic cracking aiming at a regenerator operation at much lower temperature, between 480° C. and 620° C., preferably the temperature should be between 500° C. and 600° C. The coked catalyst generated by the cracking of light streams with low potential for delta coke generation can have the combustion effected at a lower temperature. The regeneration temperature must be at least 40° C. and at most 100° C. higher than the reaction temperature, keeping the catalyst circulation high to maintain the energy balance in the reaction section. The minimum regeneration temperature can be ensured by installing an air preheating furnace before entering the regenerator and passing through the air distributor inside the regenerator. The used catalyst must contain zeolite with pores of intermediate size. Such conditions greatly favor the production of aromatics and the octane rating of the produced naphtha.

Claims

1-A PROCESS FOR OBTAINING AROMATICS from streams with linear chains with 5 to 18 carbon atoms, characterized in that it comprises the following steps: a) introducing the preheated feed stream (2) into a fluid catalytic cracking reactor (5) so that it comes into contact with an intermediate pore zeolite catalyst (3) from a regenerator (12), from the which a mixture between gaseous cracking products and coke-deactivated catalyst (6) is obtained; b) separating from said mixture the cracking products and the coke-deactivated catalyst (6); c) subjecting the coke-deactivated catalyst (6), separated in the previous step, to a rectification step (10) with water vapor (9) for removing light hydrocarbons (8); d) subjecting the rectified catalyst (11) in the previous step to a regeneration step (12) and combusting the coke deposited on the catalyst particles, in which the combustion temperature does not exceed 600° C., in order to obtain catalyst particles with activity higher than that of the spent catalyst; e) allocating the regenerated catalyst with restored activity back to the reactor (5) in order to continue the catalytic cracking process; f) allocating the hydrocarbon streams (7) obtained in the catalytic cracking reactor to a fractionation section in order to separate the produced hydrocarbons according to their boiling points; wherein the catalyst regeneration temperature is at least 40° C. and at most 100° C. higher than the reaction temperature.

2-THE PROCESS according to claim 1, characterized in that the pre-heated feed stream (2) contains triglycerides with carbonic chain fatty acids (C9-C18) from plant and/or animal biomass.

3-THE PROCESS according to claim 1, characterized in that the preheated feed stream (2) contains free carbonic chain fatty acids (C9-C18).

4-THE PROCESS according to claim 1, characterized in that the preheated feed stream (2) contains triglycerides with concentrations greater than 65% by mass of fatty acids.

5-THE PROCESS according to claim 2, characterized in that the triglyceride is selected from soybean oil (Glycine max), castor oil (Ricinus communis), cottonseed oil (Gossypium hirsutum or G. barbadenseis), palm oil (Elaies guinensis), pine oil (Tall oil), sunflower oil (Helianthus annuus), jatropha oil (Jatropha Curcas), algae oil, beef tallow (Tallow) and other oils of animal or plant origin or any other triglyceride or reaction product of its transesterification with methanol or ethanol.

6-THE PROCESS according to claim 1, characterized in that the preheated feed stream (2) includes a single oil or a mixture of two or more oils, in any ratios.

7-THE PROCESS according to claim 1, characterized in that the preheated feed stream (2) is a stream of fossil hydrocarbons selected from oil refining streams that have boiling points in the range between 40° C. and 300° C.

8-THE PROCESS according to claim 1, characterized in that the intermediate pore zeolitic catalyst (3) is selected from the types of structure consisting of: MFI, MEL, ZSM-8, ZSM-12, ZSM-21, ZSM-23, ZSM-35, ZSM-38, IMF and TUN and any combination thereof.

9-THE PROCESS according to claim 1, characterized in that the reactor (5) operates at pressures in the range of 200 to 400 kPa, catalyst-oil ratio in the range of 5 to 30, and contact time in the range between 1.0 and 3.0 seconds.

10-THE PROCESS according to claim 1, characterized in that the hydrocarbon streams (7) obtained in the catalytic cracking reactor (5) are fuel gas, liquefied gas, naphtha and other hydrocarbons.

11-THE PROCESS according to claim 1, characterized in that the fluid catalytic cracking reactor (5) is of the riser type.

12-THE PROCESS according to claim 1, characterized in that the air (13) used in the combustion of the coke deposited on the catalyst is preheated through a furnace (16) to complement the thermal balance, being fed into the regeneration section (12) at temperatures from 400° C. to 550° C.

13-AN AROMATIC STREAM as obtained in the process defined in claim 1, characterized in that it has a typical naphtha distillation range between 30° C. and 220° C.

14-THE AROMATIC STREAM according to claim 13, characterized in that it is 100% of renewable origin.

15-THE AROMATIC STREAM according to claim 14, characterized in that it contains between 70% m/m and 90% m/m of aromatics with a benzene ring in mixture with paraffinic and olefinic compounds.

16-THE AROMATIC STREAM according to claim 14, characterized in that it has at least 100 RON octane rating and 85 MON octane rating.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0038] The present invention will be described in more detail below, with reference to the attached figures which, in a schematic way and not limiting the inventive scope, represent examples of the embodiment of the same. In the drawings, there are:

[0039] FIG. 1 illustrating the structure of a typical triglyceride;

[0040] FIG. 2 illustrating in a simplified way the flow diagram of the process of the present invention;

[0041] FIG. 3 illustrating a graph with the operating conditions, reaction temperature (TRX) and dense phase temperature (TFD), used in Examples 2, 3 and 4. The catalyst-oil ratio remains constant for the same delta temperature of the dense phase of the regenerator and reaction temperature. The delta TFD-TRX was maintained at 80° C., keeping the catalyst circulation constant;

[0042] FIG. 4 illustrating a graph of the closure of the mass balance as a function of the reaction temperature;

[0043] FIG. 5 illustrating (a) conversion and (b) fuel gas as a function of reaction temperature (TRX);

[0044] FIG. 6 illustrating (a) LPG and (b) propylene as a function of reaction temperature (TRX);

[0045] FIG. 7 illustrating Naphtha as a function of reaction temperature (TRX);

[0046] FIG. 8 illustrating (a) LCO and (b) decanted oil as a function of reaction temperature (TRX);

[0047] FIG. 9 illustrating (a) water and (b) coke as a function of reaction temperature (TRX);

[0048] FIG. 10 illustrating (a) carbon dioxide and (b) carbon monoxide (as reactor products) as a function of reaction temperature (TRX);

[0049] FIG. 11 illustrating (a) CO.sub.2/CO ratio and (b) carbon in the regenerated catalyst (RCC) as a function of dense phase temperature (TFD).

DETAILED DESCRIPTION OF THE INVENTION

[0050] The present invention discloses a process for the production of aromatics from streams containing linear chains with 5 to 18 carbon atoms, of fossil or renewable origin, with application in the field of catalytic cracking through operating the regenerator at a temperature much lower temperature, between 480° C. and 620° C., preferably the temperature should be between 500° C. and 600° C., and in which the catalyst regeneration temperature is at least 40° C. and at most 100° C. higher than the reaction temperature.

[0051] A lower temperature in the regenerator promotes a lower thermal shock, but is still able to keep a high catalyst circulation to maintain the energy balance in the reaction section, provided that the reaction temperature remains at adequate values, according to Equation 1.

[0052] The coked catalyst generated by the cracking of light streams with low potential for delta coke generation can have the combustion effected at a lower temperature.

[0053] The minimum combustion temperature can be ensured by installing an air preheating furnace before entering the regenerator and passing through the air distributor inside the regenerator. The furnace must be designed for heating the air to a temperature of 550° C.

[0054] The process of the invention for obtaining aromatics and high-octane-rating gasoline from streams with linear chains of 5 to 18 carbon atoms comprises the following steps, described with the aid of FIG. 2, which presents, in a simplified way, the process flow diagram: [0055] a) introducing the preheated feed stream (2) into a fluid catalytic cracking reactor (5) so that it comes into contact with an intermediate pore zeolite catalyst (3), coming from a regenerator (12), from the which a mixture between gaseous cracking products and coke-deactivated catalyst (6) is obtained; [0056] b) separating from said mixture the cracking products and the coke-deactivated catalyst (6); [0057] c) subjecting the coke-deactivated catalyst (6), separated in the previous step, to a rectification step (10) with water vapor (9) for removing light hydrocarbons (8); [0058] d) subjecting the rectified catalyst (11) in the previous step to a regeneration step (12) and combusting the coke deposited on the catalyst particles, in which the combustion temperature does not exceed 600° C., in order to obtain catalyst particles with activity higher than that of the spent catalyst; [0059] e) allocating the regenerated catalyst (3) and with restored activity back to the reactor (5) in order to continue the catalytic cracking process; [0060] f) allocating the hydrocarbon streams (7) obtained in the catalytic cracking reactor to a fractionation section in order to separate the produced hydrocarbons according to their boiling points; [0061] wherein the catalyst regeneration temperature is at least 40° C. and at most 100° C. higher than the reaction temperature.

[0062] Next, the invention is described in detail with the help of FIG. 2, which presents, in a simplified way, the process flow diagram in a catalytic cracking unit for processing streams containing linear chains with 5 to 18 carbon atoms carbon, which mainly aims at maximizing the production of aromatics with a benzene ring.

[0063] A preheated feed stream (2) is introduced to the fluid catalytic cracking reactor (5) with the aid of a dispersing fluid such as water vapor (4), coming into contact with an intermediate pore zeolite catalyst (3), coming from the regenerator (12). The catalyst is lifted towards the reaction zone by means of lift vapor (1). Feed and catalyst come into contact under conditions of catalytic cracking, namely: temperature from 40° C. to 100° C. lower than that of the dense phase of the regeneration section (12); contact time in the range between 1.0 and 3.0 seconds; catalyst-oil ratio between 5 and 30.

[0064] At the outlet of the reactor, at the end of the reactions, the coke-deactivated catalyst (6) is separated from the products of the cracking reactions.

[0065] The coke-deactivated catalyst (6) goes to a rectification step (10) where it receives a gaseous stream (9), inert, preferably water vapor, to remove light hydrocarbons (8) that are directed to mix with the already separated products and compose hydrocarbon streams (7) obtained in the process.

[0066] After the rectification (10), a deactivated catalyst (11) goes to a regeneration step (12) by combustion of coke in the presence of air (13), heated in a preheating furnace (16) and fed by air (15) coming from a blower (not shown in FIG. 2), resulting in combustion gases (14) whose main components are carbon monoxide, carbon dioxide, nitrogen and unreacted oxygen.

[0067] The regenerated catalyst returns to the reactor (5) at a temperature high enough to provide heat for the endothermic reactions of the process, thus completing a cycle of the process of the present invention.

[0068] The hydrocarbon streams (7) obtained in the process comprise: fuel gas (hydrogen, C1 and C2) which includes ethylene, lights (C3 and C4); highly aromatic naphtha (C5+, 220° C.); and other hydrocarbons (>220° C.).

[0069] The preheated feed stream (2) of the present invention can be a biomass of plant or animal origin composed of mono, di, or triglycerides, wherein the oil of renewable origin can be castor oil, soybean oil, cottonseed oil, beef tallow or even any other triglyceride or reaction product of its transesterification with methanol or ethanol.

[0070] The preheated feed stream (2) of the present invention can further be a stream of fossil hydrocarbons to be selected among oil refining streams that have boiling points in the range between 40° C. and 300° C.

[0071] The intermediate pore zeolite catalyst (3) contains a crystalline structure selected from: MFI, MEL, ZSM-8, ZSM-12, ZSM-21, ZSM-23, ZSM-35, ZSM-38, IMF and TUN and any of these combinations.

[0072] The operating conditions in the reactor (5) occur at pressures in the range of 200 to 400 kPa, catalyst-oil ratio in the range of 5 to 30 and contact time in the range between 1.0 and 3.0 seconds. The fluid catalytic cracking reactor is of the riser type.

[0073] The gasoline produced can be subsequently hydrogenated at temperatures between 350° C. and 390° C., using a sulfide system as a catalyst based on CoMo/Al.sub.2O.sub.3, containing from 2.5% to 6% by mass of cobalt and about 7 to 10% by mass of molybdenum.

[0074] The air (15) used in the regenerator (12) for the combustion of the coke deposited on the catalyst must be preheated in the furnace (16) at temperatures of up to 550° C. to generate the air (13) to be fed into the regenerator, preferably up to 500° C.

[0075] The aromatic streams obtained from the described process have a typical naphtha distillation range, between 30° C. and 220° C., are 100% renewable, in which they contain between 70% m/m and 90% m/m of aromatics with a benzene ring mixed with paraffinic and olefinic compounds, and have a minimum RON octane rating of 100 and MON octane rating of 85.

EXAMPLES

[0076] For this study, the following tests were carried out, which represent examples of the present invention, where the results of processing streams containing linear chains with 5 to 18 carbon atoms for the production of aromatics are shown.

[0077] The process as described significantly increases the production of aromatics in catalytic cracking as can be demonstrated by the examples shown here, especially the production of xylenes, especially p-xylene. There are also increases in the production of toluene and benzene.

Example 1: U-144 from Six

1.1—Load

[0078] Biodiesel of plant origin, containing approximately 18 carbon atoms, was used as the load, where its properties are shown in Table I.

TABLE-US-00001 TABLE I Biodiesel characterization Analysis Method Result Flash point (° C.) ASTM D93/ASTM D93 150.0/174.5 Density 20/4° C. ASTM D4052/ASTM 0.8800/0.8772 Appearance/color Visual LII/brown Acidity number (mg) EN ISO 14104 0.35 Cold plugging point NBR 14747 1 Oxidation stability EN 14112 9.0 KF water content EN ISO 12937 340 (mg/kg) Viscosity @40° C. (cSt) ASTM D445 4.2559 Viscosity @60° C. (cSt) ASTM D445 2.9466 Ash content (wt. %) ASTM D482 0.001 Asphaltenes with n- ASTM D6560 <0.50 heptane Basic nitrogen (mg/kg) UOP 269-10 1.9 Higher calorific value ASTM D240 9557 Micro Carbon Residue ASTM D4530 <0.01 Aniline point (° C.) NBR 11343 <0 Sulfur (mg/kg) ASTM D5453 <5 Total nitrogen (mg/kg) ASTM D5453 17.30 Fe (mg/kg) ASTM D5708 0.19 Ni/V (mg/kg) ASTM D5708 <0.01/<0.01 Na/Zn (mg/kg) ASTM D5708 0.08/0.01

[0079] The radicals R1, R2 and R3 of the plant oil in FIG. 1 that generate biodiesel are linear chains and often have a double bond and therefore do not contain aromatic rings.

1.2—Catalyst

[0080] In this example 1, a catalytic system consisting of a mixture of a catalyst containing Y zeolite (commercial catalyst C1) and catalyst based on ZSM-5 zeolite (catalyst C2) was used. The catalyst containing ZSM-5 was then mixed with the commercial catalyst containing Y zeolite in the ratio 80%/20%, generating catalyst C4.

[0081] Catalyst C2 containing fresh ZSM-5 did not undergo prior deactivation, as catalysts based on ZSM-5 are very resistant to deactivation. The commercial catalyst with Y zeolite (catalyst C1) was collected directly from the refinery.

[0082] Catalyst C3 is also a commercial catalyst containing only Y zeolite. The characterization of the catalysts used is shown in Table II.

TABLE-US-00002 TABLE II Characterization of the catalysts C1 C2 C3 Al.sub.2O.sub.3 (wt. %) 36.2 23.0 43.2 Na (wt. %) 0.3 0.3 0.23 SiO.sub.2 (wt. %) 58.4 63.1 — RE.sub.2O.sub.3 (wt. %) 0.87 0.15 2.59 P.sub.2O.sub.5 (wt. %) 1.5 11.2 0.7 Ni (ppm) 3000 — 1053 V (ppm) 600 — 544 Surface area (m.sup.2/g) 145.4 76.0 159.2 Mesopores area (m.sup.2/g) 24 — — Micropore volume (cm.sup.3/g) 0.057 0.027 — LOI (wt. %) — 10.5 — TMP (micrometers) 83 79 — <40 micrometers (%) 0 13 —

1.3—Test Units

[0083] The U-144, located in São Mateus do Sul-PR, is a demonstration-scale FCC circulating unit, which normally operates with load flow rates of 200 kg/h, in operation with heavy vacuum gas oil (GOP), or 90 kg/h, in operation with atmospheric residue (RAT). The unit catalyst inventory is extremely high, around 300 kg. The unit has an adiabatic temperature control system for the main equipment: riser reactor, rectifier, and regenerator, which allows studies to be carried out involving the energy aspects of the process.

[0084] The cracking of triglycerides or biodiesel generates a low coke yield, leading the FCC to an energy deficit in its thermal balance. In this way, the regenerator was kept heated by preheating the air to 500° C.

[0085] Biodiesel was fed into the base of the riser reactor.

1.4—Analyses

[0086] The following yield groups were defined: fuel gas (methane, hydrogen, ethane and ethylene), LPG (C3 and C4 hydrocarbons, except propylene), propylene, naphtha (C5-220° C.), LCO (220-343° C.), decanted oil (OD: +343° C.), coke, carbon monoxide, carbon dioxide and water.

[0087] The coke yield was calculated from the combustion gas mass flow rate and its chromatographic composition. Samples of the total liquid effluent were collected to carry out the simulated distillation (ASTM D2887). For a detailed characterization of the naphtha fraction, the liquid product was further subjected to the PIANO method, which provides the distribution of hydrocarbons (n-paraffins, i-paraffins, aromatics, naphthenics, olefins) with a boiling point of up to 220° C. in mass base. The MON (Motor Octane Number) and RON (Research Octane Number) octane ratings were also calculated from the gas chromatography.

[0088] In general, studies in the literature calculate the water produced by the difference between 100% by weight and the sum of the other yields or simply do not report how the calculation of the water yield was performed. In this example, the water yield was estimated by totaling the water calculated at the end of each experiment subtracted from the value fed into the unit (lifting, dispersion, rectification and separation devices).

[0089] The catalyst-oil ratio (CTO) was calculated by the coke yield divided by the difference between carbon content in the spent and regenerated catalyst.

1.5—Experimental Result

[0090] In the state-of-the-art experiments (T1 and T2), the regeneration dense phase temperature was maintained at 690° C. by burning torch oil. Test T1 uses a reaction temperature (TRX) of 520° C., load temperature (TC) of 30° C., and conventional catalyst based on faujasite C3, while test T2 uses the same TRX of T1, but with the catalyst C4. In the experiments related to this invention (A1 to A4), the dense phase temperature was reduced to 600° C., aiming at maximizing the catalyst-oil ratio (CTO) and catalytic reactions to the detriment of thermal reactions. Catalyst C4 containing ZSM-5 was used, and the difference between regeneration temperature and reaction temperature ranged from 80° C. to 100° C.

[0091] The pressure in the riser was maintained at 257 kPa. The rectification temperature was approximately 20° C. lower than the reaction temperature in all cases.

[0092] The results in Table III show the product yields and the mass concentrations of olefins, naphthenics and aromatics of interest calculated from the PIANO method in the naphtha range.

[0093] High conversions were achieved in all cases, between 87.4% by weight and 89.0% by weight (at 520° C.), that is, a minimal difference between the highest and lowest conversion.

TABLE-US-00003 TABLE III Operating conditions on the U-144, yields (% by weight in relation to the fed load) , concentrations of olefins, naphthenic-olefins and/or diolefins, aromatics and octane rating in the naphtha Test A1 A2 A3 A4 T1 T2 Catalyst C4 C4 C4 C4 C3 C4 TRX (° C.) 500 500 510 520 520 540 TFD (° C.) 600 600 600 600 690 690 TC (° C.) 100 150 100 100 30 200 Load flow rate (kg/h) 130 130 130 130 200 180 Yields (wt. %) Fuel gas 4.0 4.3 4.8 6.0 2.9 5.3 LPG (ex-C3═) 14.3 14.4 16.1 14.9 8.4 13.1 Propylene 11.9 12.4 13.6 15.1 1.1 15.0 Naphtha, C5-220° C. 43.8 42.3 40.1 37.1 42.5 39.3 benzene — — — 1.7 0.7 1.6 toluene — — — 7.9 2.7 6.8 o-xylene — — — 0.5 0.3 0.4 m-xylene — — — 0.4 1.0 0.4 p-xylene — — — 8.1 0.4 6.9 ethylbenzene — — — 2.2 0.7 1.8 Aromatics C6-C8 — — — 20.8 5.8 17.9 LCO, 220° C.-344° C. 9.1 9.3 8.7 8.4 25.4 8.5 OD, +344° C. 3.1 3.3 2.8 2.6 3.5 2.7 Water 8.5 8.5 8.1 8.4 10.6 9.5 Coke 1.2 1.4 1.3 1.3 3.0 1.5 CO and CO.sub.2 4.1 4.1 4.6 5.9 2.6 5.2 Total 100.0 100.0 100.0 100.0 100.0 100.0 Naphtha concentration Olefins 17.0 16.5 14.2 12.8 42.5 19.6 Naphthenics/Diolefins 6.7 6.4 5.8 5.4 9.7 9.4 Aromatics 67.3 67.8 71.7 73.7 13.8 62.4 MON 96.6 96.8 98.1 98.6 77.2 94.0 RON 101.8 102.1 103.5 104.3 92.4 100.0 CTO 17.5 16.7 21.0 29.3 6.9 12.1

[0094] The concentration of naphthenic-olefins and/or diolefins provides a good indication of the stability of naphtha (directly proportional to potential gum): the lower the concentration, the greater the stability of gasoline. These concentrations are much lower in tests A1 to A4 with reduced regeneration temperature (TFD) compared to T2, that is, confirming the positive effect of the reduction in TFD, and the consequent increase in catalyst circulation (catalyst-oil ratio), on the stability of naphtha. The highest concentration of aromatics and naphtha stability was achieved at a reaction temperature of 520° C. (test A4).

[0095] In addition to stability, the benefit of the invention in relation to octane rating is clear. The octane ratings, MON and RON, are much higher in tests A1 to A4 with regeneration temperature (TFD), reduced in relation to T2, that is, confirming the positive effect of the reduction of TFD, and the consequent increase in catalyst circulation (catalyst-oil ratio), on naphtha octane rating. The highest octane rating was reached at the reaction temperature of 520° C. (test A4).

[0096] Table III also shows the percentage yields of o-xylene, m-xylene, p-xylene and ethylbenzene in relation to the load, including the state of the art T1. The total yield of aromatics with 6 to 8 carbon atoms achieved using catalysts C4 (with ZSM-5 in the test T2) was much higher than that obtained with the conventional catalyst C3 in the test T1 (17.9% by weight vs. 5.8% by weight). However, the increase in catalyst circulation obtained by applying the invention raises the aromatics again, reaching a total of 20.8% by weight in the test A4, with p-xylene being the aromatic that underwent the greatest increase, followed by toluene and of ethylbenzene. The other aromatics remained practically unchanged or suffered some decrease in their yields.

[0097] Among the C8 aromatics, p-xylene was the main compound formed when the ZSM-5 zeolite predominated in the catalytic system, while m-xylene was the predominant species with Y zeolite (C3), probably due to the smaller pore size of ZSM-5.

Example 2: FCC CC3 Pilot Plant with Biodiesel

2.1—Load

[0098] The same biodiesel from EXAMPLE 1 was used as the load, whose properties are shown in Table I.

2.2—Catalyst

[0099] The catalyst C5 was used, based on ZSM-5, containing only the ZSM-5 zeolite, and therefore without the Y zeolite (faujasite) that predominates commercially in catalytic cracking units. The fresh catalyst underwent prior deactivation at 800° C. for 5 hours, 100% vapor (2 g.Math.min.sup.−1). The characterization of the used catalyst is shown in Table IV.

TABLE-US-00004 TABLE IV Characterization of the catalyst C5 based on ZSM-5. C5 Surface area (m.sup.2/g) 149.6 Mesopores area (m.sup.2/g) 54.0 Micropore volume (cm.sup.3/g) 0.042

2.3—Test Units

[0100] The CC3 pilot unit is an FCC circulating unit that normally operates with load flow rates between 200 g/h and 1 kg/h. The unit catalyst inventory is around 2 kg. The unit has an adiabatic temperature control system in the riser, which allows studies involving the energy aspects of the process to be carried out.

[0101] The cracking of biodiesel decreases the coke yield, which would lead the FCC unit to an energy deficit in an industrial converter. However, in CC3, the regenerator is heated to the desired temperature only by using resistors.

2.4—Analyses

[0102] The following yield groups were defined: fuel gas (methane, hydrogen, ethane and ethylene), LPG (C3 and C4 hydrocarbons, except propylene), propylene, naphtha (C5-220° C.), LCO (220-343° C.), decanted oil (OD: +343° C.), coke, carbon monoxide, carbon dioxide and water.

[0103] The coke yield was calculated from the combustion gas mass flow rate and its chromatographic composition. Samples of the total liquid effluent were collected to carry out the simulated distillation (ASTM D2887). For a detailed characterization of the naphtha fraction, the liquid product was further subjected to the PIANO method, which provides the distribution of hydrocarbons (n-paraffins, i-paraffins, aromatics, naphthenics, olefins) with a boiling point of up to 220° C. in mass base. The MON and RON octane ratings were further calculated from the gas chromatography.

[0104] In general, studies in the literature calculate the water produced by the difference between 100% by weight and the sum of the other yields or simply do not report how the calculation of the water yield was performed. In this study, the water yield was estimated by totaling the water calculated at the end of each experiment subtracted from the value fed into the unit.

[0105] The catalyst-oil ratio (CTO) was calculated by the heat exchanger system with air existing in CC3 located in the catalyst transfer line.

2.5—Experimental Result

[0106] Table V presents the results obtained with biodiesel. Some yields can also be seen in FIGS. 5 to 11. The regeneration and reaction temperatures were gradually reduced, always keeping the delta of these temperatures at 80° C.

[0107] The pressure in the riser was maintained at 273 kPa.

[0108] The mass balances of the unit ended between 93.4% by weight and 101.2% by weight as shown in FIG. 4.

[0109] The results in Table V show the product yields and the mass concentrations of olefins, naphthenics and aromatics of interest calculated from the PIANO method. High conversions were achieved in all cases, between 87.8% by weight and 93.9% by weight (at 500° C.), i.e., a minimal difference between the highest and lowest conversion. The concentrations of aromatics were even higher, in tests B1 to B4, than those obtained by the results of the state of the art in tests T1 and T2.

[0110] The concentration of naphthenic-olefins and/or diolefins provides a good indication of the stability of naphtha: the lower the concentration, the greater the stability of gasoline. Naphthenic-olefins reach a minimum value at a reaction temperature of 500° C., increasing at lower temperatures. The octane rating follows the aromatics content, reaching its highest value also at 500° C., 88.9 and 103.4 for MON and RON, respectively, in the test B2. Thus, the highest concentration of aromatics, highest octane rating and highest naphtha stability were achieved at a reaction temperature of 500° C. and a regeneration temperature of 580° C.

[0111] Another relevant point is the selectivity obtained among the produced aromatics. The percentage yields of o-xylene, m-xylene, p-xylene and ethylbenzene in relation to the load are much higher than those achieved by the state of the art T1, obtaining a further increase in the yield of aromatics (25.5 wt. % in the test B2 against 17.8 wt. % in the test T2, the best result in the state of the art), with p-xylene having the highest increase in absolute value.

[0112] Among the C8 aromatics, p-xylene was the main compound formed when the ZSM-5 zeolite predominated in the catalytic system, while m-xylene was the predominant species with Y zeolite (test T2), probably due to the smaller pore size than ZSM-5.

TABLE-US-00005 TABLE V Operating conditions in CC3, yields (% by weight), concentrations of olefins, naphthenic-olefins and/or diolefins, aromatics and octane rating in the naphtha produced in the reaction with biodiesel. Test B1 B2 B3 B4 B5 B6 Catalyst C5 C5 C5 C5 C5 C5 TRX (° C.) 520 500 480 460 440 420 TFD (° C.) 600 580 560 540 520 500 TC (° C.) 100 100 100 100 100 100 Load flow rate (g/h) 375 368 380 366 366 377 Yields (wt. %) Fuel gas 5.6 5.2 4.3 3.6 3.2 2.6 LPG (ex-C3═) 16.7 17.9 15.6 15.7 13.1 13.1 Propylene 13.7 13.9 11.3 10.8 8.6 8.0 Naphtha, C5-220° C. 42.4 46.4 50.7 50.1 54.4 54.1 benzene 2.3 2.6 2.6 2.4 1.6 1.2 toluene 8.1 9.0 8.1 8.5 6.0 4.8 o-xylene 0.7 0.8 0.8 0.8 0.9 1.0 m-xylene 0.3 0.2 — 0.3 4.9 3.6 p-xylene 8.3 9.8 — 9.7 3.1 2.5 ethylbenzene 2.5 3.0 3.1 3.0 2.6 2.1 Aromatics C6-C8 22.2 25.4 24.5 24.9 19.1 15.0 LCO, 220° C.-344° C. 8.8 4.9 5.2 5.9 7.1 9.1 OD, +344° C. 3.3 1.2 1.4 1.7 2.2 3.0 Water 5.1 6.1 6.9 8.4 7.0 5.8 Coke 0.8 0.9 1.3 1.0 1.7 2.0 CO 3.0 3.0 2.9 2.3 2.2 1.8 CO.sub.2 0.6 0.4 0.5 0.4 0.6 0.5 Total 100.0 100.0 100.0 100.0 100.0 100.0 Naphtha concentration Olefins 15.8 14.7 16.2 16.1 23.0 29.6 Naphthenics/Diolefins 8.6 7.5 8.9 8.3 11.8 13.3 Aromatics 72.0 74.7 71.4 72.0 60.6 51.7 MON 87.6 88.9 87.7 88.1 84.6 82.7 RON 102.5 103.4 102.3 102.7 99.7 97.9 CTO 19.4 26.2 24.8 22.6 22.6 20.2

[0113] FIGS. 5 to 10 show the yields as a function of the reaction temperature obtained in CC3. The optimum reaction temperature for maximizing yields is between 460° C. and 500° C. Lower reaction temperatures led to considerable increases in the yields of LCO, decanted oil and coke, evidencing difficulties in vaporizing the load.

[0114] This series of experiments also allowed evaluating the efficiency of coke combustion under extreme conditions with very low dense phase temperatures. FIG. 11 shows the CO.sub.2/CO ratio and carbon in the regenerated catalyst (RCC) as a function of regeneration temperature (TFD). In conditions closer to conventional ones, CC3 operates in total combustion; however, a clear tendency to partial combustion can be observed as the dense phase temperature is reduced.

[0115] As for the RCC, the regeneration of the catalyst remains adequate even at lower temperatures, but it becomes quite high at 500° C., showing in this case some deficiency in combustion.

[0116] In experiments A1 to A4, T1 and T2 for the production of bio-aromatics carried out on a demonstration scale, the burning of torch oil was necessary to maintain the regeneration temperature at 600° C. However, experiments B1 to B6 demonstrated that it is possible to decrease the regeneration and reaction temperatures simultaneously without compromising the coke combustion on the catalyst, taking the dense phase temperature up to 500° C. In this way, only the regenerator air pre-heater would be required up to a temperature of 500° C. to maintain the thermal balance. The combustion of the coke formed on the catalyst would be responsible for obtaining some more temperature than the 500° C. offered by the air heater. For a coke yield of 0.8% by weight, it is calculated that we would have approximately 80° C. more at the dense phase temperature. In this way, the dense phase temperature of the regenerator would balance between 550° C. and 580° C., eliminating the burning of torch oil in the regenerator and therefore improving the economy of the process.

[0117] Thus, concentrations and yields of aromatics much higher than those obtained using the state of the art were achieved: the total aromatics with 6 to 8 carbons is 25.5% by weight and p-xylene reached 9.8% by weight in the invention, while for the state of the art these values reach only 13.5% in weight and 2.3% in weight, respectively, and even so they are not able to solve the problem of energy deficit without using torch oil burning.

Example 3: Pilot Plant of FCC CC3 with Beef Tallow

3.1—Load

[0118] Beef tallow was used as the load and its properties are shown in Table VI.

TABLE-US-00006 TABLE VI Beef tallow characterization Analysis Method Result Flash point (° C.) ASTM D92/90 250 pH — 4.0 Density 20/4° C. ASTM D4052 0.9106 Simulated Distillation ASTM D2887 IEP, ° C. ASTM D2887 353.8  5 wt. %, ° C. ASTM D2887 537.6 10 wt. %, ° C. ASTM D2887 574.8 20 wt. %, ° C. ASTM D2887 586.6 30 wt. %, ° C. ASTM D2887 591.6 50 wt. %, ° C. ASTM D2887 599.0 70 wt. %, ° C. ASTM D2887 603.0 80 wt. %, ° C. ASTM D2887 606.2 90 wt. %, ° C. ASTM D2887 608.4 95 wt. %, ° C. ASTM D2887 609.4 PEF, ° C. ASTM D2887 612.4 Oxygenates, m. % — 10.9

3.2—Catalyst

[0119] Catalyst C5, based on ZSM-5, from EXAMPLE 2 was used.

3.3—Test Units

[0120] The CC3 pilot unit was used, the same as in EXAMPLE 2.

3.4—Analyses

[0121] The same analyses and procedures as in EXAMPLE 2 were used.

3.5—Experimental Result

[0122] The results of Table VII show the yields of products obtained by the invention from beef tallow. The mass concentrations of olefins, naphthenics and aromatics in the naphtha produced were calculated using the PIANO method. High conversions, above 95% by weight, were achieved in both tests, D1 and D2. The production of aromatics with 6 to 8 carbon atoms in the D1 test was higher than those obtained by any of the tests of the state of the art, reaching 28.2 wt. %.

Example 4: FCC CC3 Pilot Plant with Naphtha

4.1—Load

[0123] Naphtha of fossil origin was used as the load, whose properties are shown in Table VIII. Naphtha has a very low octane rating, MON and RON, only 46.5 and 53.8, respectively, making it unsuitable for consumption in Otto cycle engines.

4.2—Catalyst

[0124] Catalyst C5, based on ZSM-5, from EXAMPLE 2 was used.

4.3—Test Units

[0125] The CC3 pilot unit was used, the same as in EXAMPLE 2.

4.4—Analyses

[0126] The same analyses and procedures as in EXAMPLE 2 were used.

TABLE-US-00007 TABLE VII Operational conditions in CC3, yields (% by weight), concentrations of olefins, aromatics and octane rating in the naphtha produced in the reaction with beef tallow. D1 D2 Catalyst C5 C5 TRX (° C.) 520 440 TFD (° C.) 600 560 TC (° C.) 100 100 Load flow rate (g/h) 362 369 Yields (wt. %) Fuel gas 8.0 3.2 LPG (ex-C3=) 22.6 20.6 Propylene 14.2 10.9 Naphtha, C5-220° C. 40.4 51.1 benzene 3.3 2.6 toluene 11.8 7.9 o-xylene 1.0 0.8 m-xylene 0.2 0.2 p-xylene 9.3 8.4 ethylbenzene 2.6 2.8 aromatics C6-C8 28.2 22.7 LCO, 220° C.-344° C. 2.6 3.6 OD, +344° C. 0.7 1.2 Water 3.2 4.3 Coke 1.3 1.5 CO 5.6 3.1 CO.sub.2 1.5 0.6 Total 100.0 100.0 Naphtha concentration (wt. %) Olefins 8.5 20.1 Aromatics 83.7 66.8 MON 91.7 85.7 RON 105.8 100.9

TABLE-US-00008 TABLE VIII Naphtha characterization Analysis Method Result Density 20/4° C. Calculated 0.7250 Simulated Distillation ASTM D2887 IEP, ° C. ASTM D2887 75.4  5 wt. %, ° C. ASTM D2887 88.6 10 wt. %, ° C. ASTM D2887 91.8 30 wt. %, ° C. ASTM D2887 102.8 50 wt. %, ° C. ASTM D2887 119.4 70 wt. %, ° C. ASTM D2887 128.4 80 wt. %, ° C. ASTM D2887 132.8 90 wt. %, ° C. ASTM D2887 139.2 95 wt. %, ° C. ASTM D2887 143.8 PEF, ° C. ASTM D2887 168.8 Paraffins C6-C7 Gas Chromatography 15.1 Paraffins C8 Gas Chromatography 24.0 Paraffins C9-C10 Gas Chromatography 5.2 Olefins/Naphthenics C6-C7 Gas Chromatography 18.8 Olefins/Naphthenics C8 Gas Chromatography 19.7 Olefins/Naphthenics C9-C11 Gas Chromatography 4.1 Aromatics C6-C7 Gas Chromatography 0.10 Aromatics C8 Gas Chromatography 9.29 Aromatics C9-C10 Gas Chromatography 0.43 MON Gas Chromatography 46.5 RON Gas Chromatography 53.8

4.5—Experimental Result

[0127] The results of Table IX show the yields of products obtained by the invention from naphtha. The mass concentrations of olefins, naphthenics and aromatics in the produced naphtha were calculated using the PIANO method. The production of aromatics with 6 to 8 carbon atoms in the test E1 reached 16.6 wt. %. Although lower than that achieved with renewable streams, this value is much higher than the percentage of 9.4 wt. % relative to the naphtha fed into the process. Furthermore, the octane rating of the produced naphtha is much higher than that of the naphtha fed into the process, approaching the values required by the gasoline specification.

TABLE-US-00009 TABLE IX Operational conditions in CC3 using fossil naphtha in the feed, yields (% by weight), concentrations of olefins, aromatics and octane rating in the naphtha produced in the reaction. E1 Catalyst C5 TRX (° C.) 500 TFD (° C.) 580 Load flow rate (g/h) 365 Yields (wt. %) Fuel gas 4.2 LPG (ex-C3=) 20.9 Propylene 5.4 Naphtha, C5-220° C. 68.0 benzene 2.3 toluene 4.7 aromatics C8 9.7 aromatics C6-C8 16.7 LCO, 220° C.-344° C. 0.7 OD, +344° C. 0.0 Water 0.0 Coke 0.5 CO 0.34 CO.sub.2 0.05 Total 100.0 MON 74.1 RON 82.5 CTO 15.6

[0128] The present invention paves the way for the use of raw materials of low added value such as fatty acids and animal tallow for the production of aromatic petrochemicals from renewable origin in a profitable manner.

[0129] In addition to superior yields, the technology enables the use of catalytic systems that are more sensitive to hydrothermal deactivation, but more efficient for the production of light and aromatic olefins, as it avoids high catalyst regeneration temperatures.

[0130] It should be noted that, although the present invention has been described in relation to the attached drawings, it may undergo modifications and adaptations by technicians skilled on the subject, depending on the specific situation, but provided that within the inventive scope defined herein.