Process for obtaining aromatic compounds and aromatic hydrocarbon stream

Abstract

The invention relates to a process for the production of aromatic compounds from linear streams of fossil or renewable origin comprising from 5 to 18 carbon atoms, and use in the field of catalytic cracking aimed at regenerator operation at much lower temperatures of 480°C to 620°C, preferably at temperatures of 500°C to 600°C. The coked catalyst resulting from the cracking of light streams with low delta coke generation potential can be subjected to combustion at lower temperatures. The regeneration temperature should be at least 40°C higher than the reaction temperature and not more than 100°C, in order to maintain a high catalyst circulation to maintain the energy balance in the reaction section. The minimum regeneration temperature can be ensured by installing a furnace for preheating the air before it enters the regenerator and passes through the air distributors inside the regenerator. The catalyst used should comprise a zeolite with pores of medium size. Such conditions significantly favor the production of aromatic compounds and the octane number of the produced naphtha.

Description

Technical Field

[0001] This invention relates to a method for producing compounds with a single aromatic ring from a stream containing a straight chain of 5 to 18 carbon atoms from fossil or renewable sources, and its application in the field of catalytic cracking by operating a regenerator at a much lower temperature of 480°C to 620°C, preferably 500°C to 600°C, and wherein the catalyst regeneration temperature is at least 40°C and at most 100°C higher than the reaction temperature. The catalyst used must contain a molecular sieve with medium-sized pores. Background Technology

[0002] In catalytic cracking in a fluidized bed, hydrocarbon cracking occurs by contacting the load with the catalyst in a dynamic flow state within an upflow tubular reactor (also known as a riser), thereby converting the load into a lighter and more economical hydrocarbon stream. Thus, hydrocarbon streams from refineries with boiling points of 350°C to 550°C are converted into lighter hydrocarbons, primarily gasoline components with distillation ranges of 35°C to 220°C. Alternatively, liquid streams from renewable sources can be used. The reaction temperature (TRX) in the reactor is approximately 540°C at the reactor outlet.

[0003] Cracking is an endothermic reaction, and the increasing severity of the reaction makes it difficult to meet this energy requirement. In conventional catalytic cracking, this heat requirement is met by burning coke deposited on the catalyst in the reaction section. In the regeneration section, the catalyst is burned with air at a temperature of approximately 700°C. In this way, its catalytic activity is restored, and the heated catalyst can be returned to the reaction section to provide the necessary heat for the endothermic reaction.

[0004] The problem of meeting heat demand is exacerbated if the streams used as loads consist of light hydrocarbons (e.g., light-chain alkanes with 5 to 18 carbon atoms from petroleum distillation, or smaller chains from associated gas) that tend to deposit less coke on the catalyst compared to conventional heavy loads such as vacuum distillate or atmospheric residue. These light streams cannot be used directly as fuel and have low commercial value. However, they can be converted into higher value-added products such as BTX (benzene, toluene, and xylene) aromatic compounds.

[0005] Like alkyl fossil streams, some renewable liquid streams (such as soybean oil, castor oil, and tallow) have straight chains with up to 18 carbons in their structure, such as... Figure 1 As shown, a typical triglyceride structure is illustrated. The groups R1, R2, and R3 (straight carbon chains) typically have 9 to 18 carbon atoms. These streams generally contain at least one degree of unsaturation, but like fossil streams obtained from associated gas production, they do not have aromatic rings in their structure and have low coke-forming potential in catalytic cracking.

[0006] Useful catalysts for catalytic cracking include USY and REY type Y zeolites. In commercial catalytic cracking reactors, ZSM-5 (MFI) zeolite with medium-sized pores is further used to produce olefins, aromatic compounds, and to increase the octane number of gasoline. The term "mesoporous zeolite" applies to any zeolite whose pore size is between that of small-pore zeolites (e.g., type A zeolites) and macroporous zeolites (e.g., X and Y zeolites). Medium-sized zeolites are those with a crystal structure in which the rings have 10 or 12 oxygen atoms. The silicon to aluminum ratio of these zeolites must be between 10 and 300. Examples of mesoporous zeolites besides ZSM-5 include ZSM-8, ZSM-11 (also known as MEL), ZSM-12, ZSM-21, ZSM-23, ZSM-35, ZSM-38, IMF (also known as IM-5), TUN (also known as TNU-9), and EUO zeolites. Of these, only ZSM-5 is currently commercially used for catalytic cracking because many of them are unstable and lose their activity at the high temperatures currently operating in these converters.

[0007] Processing straight-chain streams with 5 to 18 carbon atoms in catalytic cracking results in a failure to meet the thermal requirements of catalytic cracking, mainly for the following three reasons: 1) The support itself is not a good precursor for coke formation, exhibiting low Langevin residue; 2) Light hydrocarbons are more difficult to crack, and existing technologies indicate the application of higher reaction temperatures; 3) ZSM-5 zeolite and other zeolites with smaller cavities than Y zeolite cannot produce large amounts of coke deposits on the catalyst.

[0008] However, there is little literature on addressing the energy deficiency problem in catalytic cracking. Typically, this deficiency is addressed by burning ignition oil (also known as heating oil) in the regenerator. In this method, the regenerator becomes a combustor, where the oil combustion generates sufficient heat to heat the catalyst. The regenerator bed must be heated to approximately 700°C by burning the ignition oil. A dense phase temperature (TFD) below 680°C makes it difficult to burn the heating oil in the regenerator bed, leading to uncontrolled catalyst circulation toward the cracking reactor. The generated heat is transferred to the reaction section via the catalyst itself. However, burning oil in the regenerator bed promotes several regenerator operation problems. For example, the choice of heating oil used must be careful, as heating oils with very low distillation points can cause post-burning, i.e., combustion outside the bed. The temperature difference between the bed and the combustion gases can reach 300°C, resulting in high temperatures in the cyclone separators and other equipment both inside and outside the regenerator. Another potential problem is wear and tear on the disperser responsible for introducing the heating oil into the regenerator. The high temperatures generated in the combustion bed can lead to further premature catalyst deactivation. Furthermore, the burning of oil results in energy waste, significantly increasing the carbon footprint of the catalytic cracking unit.

[0009] When people want to maximize the production of aromatic compounds, in addition to insufficient energy, the catalytic cracking of lipid streams (triglycerides) leads to another problem: causing a large amount of gas in the range of LPG (liquefied petroleum gas) as a reaction byproduct, because the characteristics of the molecules that make up vegetable oils and animal fats, which have long straight chains, no aromatic rings, and some double bonds in the side chains, are suitable for the production of these gases.

[0010] In this way, small refineries can greatly benefit from processes specifically designed for the production of aromatic compounds without increasing propylene or LPG yields. For these refineries, innovative methods are needed to minimize gas formation during cracking.

[0011] Gas production can be reduced by lowering the reaction temperature (TRX). This approach presents some challenges because catalytic cracking (FCC) is thermally balanced, and lowering the reaction temperature reduces the energy demand of the reaction section. This results in lower catalytic coke yields and lower temperatures in the catalyst regeneration section, where the combustion of coke deposited on the catalyst during the reaction occurs. This reduction in the dense phase temperature (TFD) in turn makes coke combustion itself more difficult. Modern regenerators operate in the dense phase of the regeneration bed, where most of the combustion of coke deposited on the catalyst occurs. The lowest temperature in the regenerator is approximately 680°C, and the catalyst residence time is approximately 4 minutes, which ensures coke combustion. Typically, the maximum temperature in the dense phase is designed to be 760°C, and in operation, this temperature is maintained at a maximum of 730°C.

[0012] However, regenerators designed in the 1950s operated at much lower regeneration temperatures (around 620°C) because the catalysts used in UFCCs after World War II were silica-alumina, which were more susceptible to hydrothermal deactivation and suffered severe deactivation at higher combustion temperatures. In this way, regenerators were designed to operate at milder temperatures, compensating for the less demanding temperatures by increasing the residence time (around 10 minutes). The advent of Y zeolite changed the regenerator design criteria, and they began to be designed for higher regeneration temperatures (680°C to 720°C) and shorter residence times (4 minutes).

[0013] For simplicity, the dense phase temperature (TFD) is considered as the regeneration temperature. It is known that for the same difference between the regeneration temperature and the reaction temperature (TFD-TRX), approximately the same catalyst-oil ratio is achieved (Equation 1). That is, catalyst cycling can be maximized by reducing the difference between the regeneration temperature and the reaction temperature.

[0014] Catalyst / oil × (TFD-TRX) = constant (Equation 1)

[0015] Light streams (fossil or renewable) with boiling points in the range of 70°C to 300°C can be converted into aromatic compounds in catalytic cracking using ZSM-5 and high reaction temperatures. In this environment, the temperature of the regenerator is maintained at approximately 700°C. The regenerated catalyst is returned to the riser reactor, thus providing heat to the reaction medium. However, the heat required for cracking reactions in the riser environment limits the reaction temperature in this environment, thereby limiting the operational severity. Furthermore, the thermal shock induced by the catalyst at the bottom of the riser reactor favors the formation of light olefins (e.g., propane, propylene, butane, and other compounds in the LPG and fuel gas range) as byproducts. This thermal shock is promoted by the very high temperature (approximately 700°C) at which the catalyst is located in the regenerator.

[0016] However, a solution to this technical problem becomes necessary when these light olefins are not desirable and the goal is to maximize the production of high-octane streams of aromatic-based liquid fractions and naphtha in a circulating fluidized bed.

[0017] Document BR112013012925-5 discloses a method for producing high-concentration and high-value-added BTX aromatic compounds by using various oil fractions (including LCO (light oil produced from recovery)) obtained by catalytic cracking, which contain a large amount of highly aromatic compounds, instead of naphtha (a conventional feedstock). This method employs hydrogenation, catalytic cracking, separation, and alkyl transfer steps.

[0018] Document US6153089 discloses a method for converting undesirable olefin streams into hydrogen and petrochemical feedstocks (e.g., light olefins in the C2 to C4 range, and aromatic compounds, especially toluene and xylene). This method involves simultaneously cracking and reforming olefin-rich hydrocarbons in a circulating fluidized bed reactor-regeneration system using a catalyst composed of a dehydrogenation metal component, a selective zeolite component, and a macroporous acid component in varying ratios. The reactor temperature is in the range of 450°C to 750°C, and the WHSV is 0.1 / h to 60 / h. A specific rate reduction is achieved by installing a fluidized bed reactor at the end of the catalytic cracking riser reactor.

[0019] One approach to mitigating the energy deficit caused by using streams with low coke formation potential involves using catalysts in catalytic cracking that contain zeolites modified with dehydrogenated metals such as nickel, which promote coke formation, as taught in document WO2010023456. While mitigating the energy deficit in cracking, the proposed solutions do not eliminate it.

[0020] Over the years, numerous studies have been conducted at various scales (pilot, demonstration and industrial) on the production of bioproducts by catalytic cracking from renewable streams from different sources, in order to produce a wide range of products also derived from renewable sources.

[0021] Some renewable streams, even liquids, possess chemical properties very different from lipid streams (soybean oil, castor oil, or tallow), exhibiting high coke-forming potential. Among these, lignocellulose streams, such as pyrolysis oils or bio-oils from rapid biomass pyrolysis processes, can be mentioned. Bio-oils consist of a mixture of several oxygenated compounds (e.g., aldehydes, ketones, phenols). They have high aromatic content and high coking potential in catalytic cracking and should not be confused with lipid streams. For example, document BR102012013787-9 discloses a method for producing fuel, more specifically high-octane gasoline, from plant biomass derived from lignocellulose, which lacks straight carbon chains. Such a process involves the following steps: first, lignin depolymerization is achieved by co-treating the plant biomass with the hydrocarbon streams, introduced separately into two reaction zones of a fluidized catalytic cracking reactor under low pressure; second, the hydrocarbon streams from the gasoline fraction of the catalytic cracking reactor are hydrogenated. In contrast to the lipid streams of this invention, bio-oils do not suffer from energy deficiency in catalytic cracking because they are a source of lignocellulose and exhibit high coke-forming potential.

[0022] Document BR102016003995-9 discloses a method for obtaining biofuels, particularly high-octane gasoline, by co-processing a gaseous stream containing pyrolysis oil (bio-oil), hydrogen-rich light saturated hydrocarbons, and a fossil-based main load in an FCC process. The process utilizes two reaction stages: a first reaction stage with reaction temperatures ranging from 500°C to 800°C, and a second reaction stage with temperatures ranging from 500°C to 620°C. The bio-oil, i.e., the plant stream described in this document, does not result in energy deficits during catalytic cracking because it is a lignocellulose source and, unlike the lipid stream of this invention, possesses high coke-forming potential.

[0023] References BRPI1004598 and BRPI1102958 teach how to maximize the production of aromatic BTX (benzene, toluene, xylene) from vegetable oils using a ZSM-5-based catalyst for the unprecedented production of polymers from renewable sources such as PET and polystyrene. The achieved concentrations of aromatic compounds in the liquid product are significantly higher than those obtained using an octahedral zeolite catalyst, as achieved by BR 102014022366-5. However, gas yields remain considerably high, and the naphtha-range product yield, in which the aromatics of interest are present, never exceeds 35 wt% relative to the feed. LPG yields remain at 30.0 wt%, which is too high for a catalytic cracking unit without a propylene separator to be marketable. The energy deficit in catalytic cracking is addressed by co-treating a lignocellulose stream (with high coke-forming potential) from the rapid pyrolysis of biomass with a lipid stream (with low coke-forming potential) to restore thermal equilibrium in the converter.

[0024] Document US7540952 teaches how to co-process vegetable oil with heavy vacuum distillate in catalytic cracking to maximize the middle distillate output. As an example, it mentions using castor oil (castor oil or ricinole) as feedstock for an FCC pilot unit operating at reaction temperatures of 300°C to 400°C, significantly lower than those used in conventional FCCs. Under these conditions, products in the diesel range are maximized, but little low-octane gasoline is produced. However, this patent does not teach how to address the increased energy deficit resulting from using large quantities of vegetable oil as feedstock to the FCC converter instead of vacuum distillate, which limits the use of this renewable distillate.

[0025] Reference BR 102014022366-5 also uses renewable lipid streams for the production of cracked naphtha in a catalytic cracking reactor using an octahedral zeolite catalyst. The produced naphtha exhibits properties similar to those of fossil-derived cracked naphtha, which has a high content of bioaromatic compounds (with the advantage of low oxygenate content), and provides a high calorific value compared to alternative products such as ethers (ETBE and MTBE) and alcohols (ethanol and butanol). However, cracked naphtha obtained from soybean oil has low oxidative stability, requiring chemical treatment to ensure the stability required for commercial gasoline. This reference also does not propose solutions to the energy deficit in catalytic cracking caused by the use of vegetable oils.

[0026] WO2012 / 088546 discloses a method for producing aromatic compounds from renewable sources using a gallium-containing catalyst. More specifically, it discloses preferred embodiments relating to the conversion of oils comprising fats or other lipids derived from biomass, such as oils from naturally occurring non-vascular photosynthetic organisms and / or from genetically modified non-vascular photosynthetic organisms; rapeseed oil and other oils derived from plants such as corn, soybean, sunflower, and sorghum; and / or oils from other plant materials, seeds, fungi, bacteria, and other living organisms. The biomass conversion is not carried out in a catalytic cracking reactor.

[0027] Document US2009 / 0158637 discloses a method for producing aviation fuel from renewable feedstocks such as vegetable oils and animal fats. The method involves treating the renewable feedstock with hydrogenation and deoxygenation to provide n-chain alkanes of about 8 to about 24 carbon atoms. At least some of the n-chain alkanes are isomerized to improve cold flow characteristics. At least a portion of the wax is selectively decomposed to provide waxes that meet different aviation fuel specifications (e.g., JP-8). This method does not employ selective catalytic cracking conditions and is not designed to maximize the production of aromatic compounds.

[0028] The prior art literature does not disclose a method for producing aromatic compounds in a fluidized catalytic bed with cold regeneration, which exhibits specific properties such as those of the present invention.

[0029] The present invention proposes operating the regenerator at a much lower temperature of 480°C to 620°C, using a light load in the range of gasoline and diesel during operation, or using a load with low coke production potential or low residual carbon.

[0030] Preferably, the temperature should be between 500°C and 600°C; thus, the lower temperature in the regenerator will promote a smaller thermal shock to the load of the catalyst when it returns to the riser reactor, but will maintain a high catalyst circulation to maintain the energy balance in the reaction section. Combustion can be achieved by subjecting the coking catalyst, which has a low δ-coke formation potential, to combustion at a lower temperature through light fluid cracking. The minimum combustion temperature can be ensured by installing an air preheater before it enters the regenerator and passes through the air distributor inside the regenerator. The regeneration temperature will be maintained at least 40°C and at most 100°C higher than the reaction temperature.

[0031] Such conditions are highly favorable for the production of aromatic compounds and the octane number of naphtha produced.

[0032] Another advantage of the solution is the minimization of catalyst deactivation, as the combustion of the coke deposited on the catalyst occurs under significantly milder conditions, thus allowing catalytic activity to be maintained.

[0033] This invention is important for the production of aromatic compounds from fossil or renewable sources, with or without unsaturation, having straight carbon chains with 5 to 18 carbon atoms. These aromatic compounds can be used as basic petrochemical feedstocks.

[0034] It further allows for the conversion of streams from generally renewable sources, such as tallow, used frying oil, vegetable oils, and fatty acids, into aromatic compounds. These same aromatic compounds can be used as feedstock for addition to jet fuel (QAV). It should be noted that there are currently no other production routes for producing renewable aromatic compounds for QAV.

[0035] As described below, the present invention advantageously presents an increase in selectivity for the production of aromatic compounds, while minimizing the problems caused by insufficient energy in the process. Summary of the Invention

[0036] This invention relates to a method for producing aromatic compounds, wherein operation in a regenerator occurs at a much lower dense-phase temperature of 480°C to 620°C, preferably 500°C to 600°C, during which a stream comprising straight chains having 5 to 18 carbon atoms from fossil or renewable sources is used, exhibiting low coke production potential and low residual carbon. The lower temperature in the regenerator will promote a smaller thermal shock in the reaction section, but if the regeneration temperature (dense-phase temperature) is at least 40°C and at most 100°C higher than the reaction temperature, the catalyst circulation will remain high to maintain energy balance in the reaction section. Combustion can be achieved at lower temperatures by subjecting the coking catalyst, produced by cracking a light stream with low δ-coke production potential, to combustion. A minimum combustion temperature can be ensured by installing an air preheater before the air enters the regenerator and passes through the air distributor inside the regenerator. Such conditions are highly favorable for the production of aromatic compounds and the octane number of the naphtha produced.

[0037] This invention can be applied to any existing catalytic cracking unit, provided that minor modifications are implemented to allow for maintaining energy balance. It is recommended to operate the combustion air heater for the regenerator at a constant temperature to maintain energy balance if necessary. The heater must be designed to heat the air to a temperature of 550°C.

[0038] The catalyst used must contain molecular sieves with medium-sized pores. Attached Figure Description

[0039] The invention will now be described in more detail with reference to the accompanying drawings, which illustrate examples of embodiments of the invention in a schematic manner rather than limiting the scope of the invention. In the drawings, the following are present:

[0040] - Figure 1 The structure of a typical triglyceride is shown;

[0041] - Figure 2 A simplified flowchart of the method of the present invention is shown.

[0042] - Figure 3 The diagram shows the operating conditions used in Examples 2, 3, and 4, namely the reaction temperature (TRX) and dense phase temperature (TFD). For the same Δ regenerator, the dense phase temperature versus the reaction temperature, the catalyst-oil ratio was kept constant. ΔTFD-TRX was maintained at 80°C, thus keeping the catalyst cycle constant;

[0043] - Figure 4 The closure curve of the mass balance as a function of reaction temperature is shown.

[0044] - Figure 5 (a) Conversion and (b) Fuel gas as a function of reaction temperature (TRX) are shown;

[0045] - Figure 6 (a) LPG and (b) propylene are shown as a function of reaction temperature (TRX);

[0046] - Figure 7 Naphtha is shown as a function of reaction temperature (TRX);

[0047] - Figure 8 (a) LCO and (b) decanted oil are shown as a function of reaction temperature (TRX);

[0048] - Figure 9 (a) water and (b) coke are shown as a function of reaction temperature (TRX);

[0049] - Figure 10 (a) carbon dioxide and (b) carbon monoxide (as reactor products) are shown as a function of reaction temperature (TRX);

[0050] - Figure 11 (a) CO2 / CO ratio and (b) carbon in the regenerated catalyst as a function of dense phase temperature (TFD) are shown. Detailed Implementation

[0051] The present invention discloses a method for producing aromatic compounds from fossil or renewable sources comprising straight chains having 5 to 18 carbon atoms, and its application in the field of catalytic cracking by operating a regenerator at a temperature much lower than 480°C to 620°C, preferably 500°C to 600°C, and wherein the catalyst regeneration temperature is at least 40°C and at most 100°C higher than the reaction temperature.

[0052] According to Equation 1, the lower temperature in the regenerator promotes a lower thermal shock, but it is still possible to maintain a high catalyst cycle to keep the energy balance of the reaction section, provided that the reaction temperature is kept at an appropriate value.

[0053] Combustion can be achieved at lower temperatures by subjecting coking catalysts produced by light fluid cracking with low δ coke generation potential to combustion.

[0054] The minimum combustion temperature can be ensured by installing an air preheating furnace before the air enters the regenerator and passes through the air distributor inside the regenerator. The furnace must be designed to heat the air to a temperature of 550°C.

[0055] The method of the present invention for obtaining aromatic compounds and high-octane gasoline from a straight-chain stream containing 5 to 18 carbon atoms includes the following steps, by means of Figure 2 It was described. Figure 2 The following process flow diagram is presented in a simplified manner:

[0056] a) A preheated feed stream (2) is introduced into a fluidized catalytic cracking reactor (5) such that the preheated feed stream comes into contact with a mesoporous zeolite catalyst (3) from a regenerator (12) to obtain a mixture of gaseous cracking products and coke-deactivated catalyst (6) from the fluidized catalytic cracking reactor (5).

[0057] b) Separate the cracking products and the coke-deactivated catalyst (6) from the mixture;

[0058] c) The catalyst (6) that deactivates the coke separated in the previous step undergoes a distillation step (10) using steam (9) to remove light hydrocarbons (8);

[0059] d) subjecting the catalyst (11) distilled in the previous step to a regeneration step (12) and burning the coke deposited on the catalyst particles, wherein the combustion temperature does not exceed 600°C, to obtain catalyst particles with higher activity than the spent catalyst.

[0060] e) Distribute the regenerated catalyst (3) with restored activity back to the reactor (5) to continue the catalytic cracking process;

[0061] f) The hydrocarbon stream (7) obtained in the catalytic cracking reactor is distributed to the fractionation section to separate the produced hydrocarbons according to the boiling point of the produced hydrocarbons;

[0062] The catalyst regeneration temperature is at least 40°C and at most 100°C higher than the reaction temperature.

[0063] Next, with the help of Figure 2 The invention will be described in detail below. Figure 2A simplified process flow diagram is presented for a catalytic cracking unit that processes streams containing straight chains with 5 to 18 carbon atoms, with the primary objective of maximizing the production of aromatic compounds with benzene rings.

[0064] The preheated feed stream (2) is introduced into the fluidized catalytic cracking reactor (5) by means of a dispersing fluid such as steam (4), thereby contacting the mesoporous zeolite catalyst (3) from the regenerator (12). The catalyst is lifted toward the reaction zone by a lifting steam (1). The feed and catalyst are contacted under catalytic cracking conditions, namely: a temperature 40°C to 100°C lower than the temperature of the dense phase in the regeneration section (12); a contact time in the range of 1.0 seconds to 3.0 seconds; and a catalyst-to-oil ratio of 5 to 30.

[0065] At the reactor outlet, at the end of the reaction, the coke-deactivated catalyst (6) is separated from the products of the cracking reaction.

[0066] The deactivated catalyst (6) from the coke enters the distillation step (10), in which it receives an inert gas stream (9), preferably steam, to remove light hydrocarbons (8), which are directed to mix with the separated products and form the hydrocarbon stream (7) obtained in the process.

[0067] After distillation (10), the deactivated catalyst (11) enters the regeneration step (12), and is then heated in the preheating furnace (16) and passed through a blower (in Figure 2 In the presence of air (13) fed by air (15, not shown), coke is burned to obtain a combustion gas (14) whose main components are carbon monoxide, carbon dioxide, nitrogen and unreacted oxygen.

[0068] The regenerated catalyst is returned to the reactor (5) at a temperature high enough to provide the heat of the endothermic reaction of the process, thus completing the cycle of the process of the present invention.

[0069] The hydrocarbon stream (7) obtained in this process includes: fuel gas containing ethylene (hydrogen, C1 and C2), light substances (C3 and C4); highly aromatic naphtha (C5+, 220°C); and other hydrocarbons (>220°C).

[0070] The preheated feed stream (2) of the present invention can be a plant or animal-derived biomass consisting of monoglycerides, diglycerides or triglycerides, wherein the oil of renewable source can be castor oil, soybean oil, cottonseed oil, tallow, or even any other triglyceride or its transesterification product with methanol or ethanol.

[0071] The preheated feed stream (2) of the present invention can also be a stream of fossil hydrocarbons selected from refining streams with boiling points in the range of 40°C to 300°C.

[0072] The mesoporous zeolite catalyst (3) comprises crystal structures selected from the following: MFI, MEL, ZSM-8, ZSM-12, ZSM-21, ZSM-23, ZSM-35, ZSM-38, IMF and TUN, and any combination thereof.

[0073] The operating conditions in reactor (5) are as follows: pressure ranging from 200 kPa to 400 kPa, catalyst-to-oil ratio ranging from 5 to 30, and contact time ranging from 1.0 seconds to 3.0 seconds. The fluidized catalytic cracking reactor is a riser type.

[0074] The resulting gasoline can then be hydrogenated at temperatures of 350°C to 390°C using a sulfide system based on a CoMo / Al2O3 catalyst (containing 2.5% to 6% cobalt and about 7% to 10% molybdenum).

[0075] The combustion air (15) used in the regenerator (12) for the coke deposited on the catalyst must be preheated in the furnace (16) at a temperature of up to 550°C to produce the air (13) to be fed into the regenerator (preferably up to 500°C).

[0076] The aromatic streams obtained by the process have a typical naphtha range of 30°C to 220°C, are 100% renewable, and contain 70% to 90% by mass of aromatic compounds with benzene rings mixed with alkane and olefin compounds, and have a minimum RON octane number of 100 and a MON octane number of 85.

[0077] Example:

[0078] For this study, the following experiments were conducted, representing embodiments of the present invention, showing the results of processing a stream containing a straight chain having 5 to 18 carbon atoms for the production of aromatic compounds.

[0079] The described process significantly increases the yield of aromatic compounds in catalytic cracking, as can be demonstrated by the examples shown herein, particularly xylene, and especially para-xylene. The yields of toluene and benzene are also increased.

[0080] Example 1: U-144 from SIX

[0081] 1.1- load:

[0082] Plant-derived biodiesel containing approximately 18 carbon atoms was used as the loading, the characteristics of which are shown in Table I.

[0083] Table I

[0084] Biodiesel characterization

[0085]

[0086] Producing biodiesel Figure 1 The groups R1, R2, and R3 in vegetable oils are straight chains and usually have double bonds and therefore do not contain aromatic rings.

[0087] 1.2- catalyst:

[0088] In this Example 1, a catalytic system consisting of a catalyst containing Y zeolite (commercial catalyst C1) and a catalyst based on ZSM-5 zeolite (catalyst C2) was used. The ZSM-5 catalyst was then mixed with the commercial catalyst containing Y zeolite at a ratio of 80% / 20% to produce catalyst C4.

[0089] Catalyst C2, containing fresh ZSM-5, does not undergo prior deactivation because ZSM-5-based catalysts are highly resistant to deactivation. Commercial catalysts with Y zeolite (catalyst C1) are collected directly from refineries.

[0090] Catalyst C3 is also a commercial catalyst containing only Y zeolite. The characterization of the catalysts used is shown in Table II.

[0091] Table II

[0092] Characterization of catalysts

[0093]

[0094] 1.3- Experimental setup:

[0095] lie in The U-144 at Mateus do Sul-PR is a demonstration-scale FCC cycle unit, typically operating at a load flow rate of 200 kg / h (when operating with heavy vacuum distillate (GOP)) or 90 kg / h (when operating with atmospheric residue (RAT)). The unit has a very high catalyst charge, approximately 300 kg. The unit features an adiabatic temperature control system for the main equipment: riser reactor, distiller, and regenerator, which allows for studies of the energy aspects of the process.

[0096] The cracking of triglycerides or biodiesel produces low coke yields, resulting in insufficient energy for the FCC in its thermal equilibrium. In this way, the regenerator is kept heated by preheating the air to 500°C.

[0097] Biodiesel is fed into the bottom of the riser reactor.

[0098] 1.4- analyze:

[0099] The following yield groups were defined: fuel gas (methane, hydrogen, ethane, and ethylene), LPG (C3 and C4 hydrocarbons, excluding propylene), propylene, naphtha (C5-220°C), LCO (220°C to 343°C), decant oil (OD: +343°C), coke, carbon monoxide, carbon dioxide, and water.

[0100] Coke yield was calculated from the mass flow rate of the combustion gases and their chromatographic composition. Samples of all liquid effluent were collected for simulated distillation (ASTM D2887). To characterize the naphtha fraction in detail, the liquid products were further subjected to the PIANO method, which provides the mass distribution of hydrocarbons (n-alkanes, isoalkanes, aromatics, cycloalkanes, and alkenes) with boiling points up to 220°C. MON (Motor Octane Number) and RON (Research Octane Number) octane numbers were also calculated by gas chromatography.

[0101] Generally, studies in the literature calculate the water produced by the difference between 100% by weight and the sum of other yields, or do not report how the water yield is calculated at all. In this embodiment, the water yield is estimated by subtracting the total water calculated at the end of each experiment from the values ​​fed into the apparatus (lifting device, dispersing device, distillation device, and separation device).

[0102] The catalyst-to-oil ratio (CTO) is calculated by dividing the coke yield by the difference in carbon content between the spent catalyst and the regenerated catalyst.

[0103] 1.5- Experimental results:

[0104] In prior art experiments (T1 and T2), the regeneration dense phase temperature was maintained at 690°C by burning ignition oil. Experiment T1 used a reaction temperature (TRX) of 520°C, a loading temperature (TC) of 30°C, and a conventional catalyst based on octahedral zeolite C3, while experiment T2 used the same TRX as T1 but with catalyst C4. In experiments related to this invention (A1 to A4), the dense phase temperature was reduced to 600°C to maximize the catalyst-to-oil ratio (CTO) and catalytic reactions detrimental to thermal reactions. Catalyst C4 containing ZSM-5 was used, and the difference between the regeneration temperature and the reaction temperature ranged from 80°C to 100°C.

[0105] The pressure in the riser was maintained at 257 kPa. In all cases, the distillation temperature was approximately 20 °C lower than the reaction temperature.

[0106] The results in Table III show the product yields and mass concentrations of the olefins, cycloalkanes, and aromatic compounds of interest, calculated using the PIANO method within the naphtha range.

[0107] High conversion rates of 87.4% to 89.0% by weight (at 520°C) were achieved in all cases, meaning the difference between the highest and lowest conversion rates was minimal.

[0108] Table III

[0109] Operating conditions for U-144, yield (wt% relative to feed load), concentrations of olefins, cycloalkane-olefins and / or dienes, aromatic compounds, and octane number in naphtha.

[0110]

[0111] The concentration of cycloalkane-olefins and / or dienes provides a good indicator of naphtha stability (proportional to potential gum content): the lower the concentration, the better the stability. These concentrations were much lower in tests A1 to A4 using a reduced regeneration temperature (TFD) compared to T2, i.e., confirming the positive effect of reduced TFD and subsequent increased catalyst cycle (catalyst-oil ratio) on naphtha stability. The highest concentrations of aromatic compounds and naphtha stability were achieved at a reaction temperature of 520°C (test A4).

[0112] In addition to stability, the octane number-related benefits of this invention are also evident. In tests A1 to A4 using a regeneration temperature (TFD) lower than T2, octane numbers MON and RON were significantly higher, demonstrating the positive effect of reduced TFD and subsequent increased catalyst cycle (catalyst-oil ratio) on naphtha octane number. The highest octane number was achieved at a reaction temperature of 520°C (Test A4).

[0113] Table III also shows the percentage yields of o-xylene, m-xylene, p-xylene, and ethylbenzene relative to their loadings, including prior art T1. The total yields of aromatic compounds with 6 to 8 carbon atoms obtained using catalyst C4 (ZSM-5 in experiment T2) were significantly higher than the total yields obtained using conventional catalyst C3 in experiment T1 (17.9 wt% vs. 5.8 wt%). However, the increased catalyst cycling obtained by applying the present invention again improved the yields of aromatic compounds, reaching a total of 20.8 wt% in experiment A4, with p-xylene experiencing the largest increase, followed by toluene and ethylbenzene. Other aromatic compounds remained virtually unchanged in yield or experienced some decrease.

[0114] In C8 aromatic compounds, p-xylene is the dominant compound when ZSM-5 zeolite is dominant in the catalytic system, while in the case of Y zeolite (C3), m-xylene is the dominant substance, possibly due to the smaller pore size of ZSM-5.

[0115] Example 2: FCC CC3 pilot plant using biodiesel

[0116] 2.1- load:

[0117] The same biodiesel as in Example 1 was used as the load, and its characteristics are shown in Table I.

[0118] 2.2- catalyst:

[0119] A ZSM-5-based catalyst, C5, containing only ZSM-5 zeolite and therefore free of the commercially dominant Y zeolite (octahedral zeolite), was used in the catalytic cracking unit. The fresh catalyst underwent initial deactivation for 5 hours at 800 °C and 100% vapor (2 g / min). The characterization of the catalyst used is shown in Table IV.

[0120] Table IV

[0121] Characterization of C5 catalyst based on ZSM-5

[0122]

[0123] 2.3- Experimental setup:

[0124] The CC3 pilot plant is an FCC circulating unit that typically operates at a load flow rate of 200 g / h to 1 kg / h. The catalyst charge is approximately 2 kg. The unit has an adiabatic temperature control system in the riser, which allows for energy-related studies of the process.

[0125] Cracking of biodiesel reduces coke yield, which leads to insufficient energy in the industrial converter of FCC units. However, in CC3, the regenerator is heated to the desired temperature simply by using a resistor.

[0126] 2.4- analyze:

[0127] The following yield groups were defined: fuel gas (methane, hydrogen, ethane, and ethylene), LPG (C3 and C4 hydrocarbons, excluding propylene), propylene, naphtha (C5-220°C), LCO (220°C to 343°C), decant oil (OD: +343°C), coke, carbon monoxide, carbon dioxide, and water.

[0128] Coke yield was calculated from the mass flow rate of the combustion gases and their chromatographic composition. Samples of all liquid effluents were collected for simulated distillation (ASTM D2887). To characterize the naphtha fraction in detail, the liquid products were further subjected to the PIANO method, which provides the mass distribution of hydrocarbons (n-alkanes, isoalkanes, aromatics, cycloalkanes, and alkenes) with boiling points up to 220 °C. MON and RON octane numbers were further calculated by gas chromatography.

[0129] Generally, studies in the literature calculate the water produced by the difference between 100% by weight and the sum of other yields, or do not report how the water yield is calculated at all. In this study, the water yield was estimated by subtracting the total water calculated at the end of each experiment from the value fed into the device.

[0130] The catalyst-to-oil ratio (CTO) is calculated using a heat exchanger system located in the catalyst delivery line that uses air present in CC3.

[0131] 2.5- Experimental results:

[0132] Table V presents the results obtained using biodiesel. Some yields can also be found in [reference needed]. Figures 5 to 11 Gradually decrease the regeneration temperature and reaction temperature, always maintaining the Δ value of these temperatures at 80℃.

[0133] The pressure in the riser pipe is maintained at 273 kPa.

[0134] like Figure 4 As shown, the mass balance of the device terminates between 93.4% by weight and 101.2% by weight.

[0135] The results in Table V show the product yields and mass concentrations of the olefins, cycloalkanes, and aromatic compounds of interest calculated according to the PIANO method. High conversions of 87.8 wt% to 93.9 wt% (at 500 °C) were achieved in all cases, i.e., the difference between the highest and lowest conversions was minimal. In experiments B1 to B4, the concentrations of aromatic compounds were even higher than those obtained using prior art results in experiments T1 and T2.

[0136] The concentration of cycloalkane-olefins and / or dienes provides a good indicator of naphtha stability: the lower the concentration, the better the stability. The concentration of cycloalkane-olefins reaches a minimum at a reaction temperature of 500°C and increases at lower temperatures. The octane number follows the aromatic compound content and also reaches its maximum at 500°C in test B2, with values ​​of 88.9 for MON and 103.4 for RON. Therefore, the highest aromatic compound concentration, the highest octane number, and the highest naphtha stability were obtained at a reaction temperature of 500°C and a regeneration temperature of 580°C.

[0137] Another relevant point is the selectivity obtained in the production of aromatic compounds. The percentage yields of o-xylene, m-xylene, p-xylene, and ethylbenzene relative to the load were significantly higher than those achieved by prior art T1, resulting in a further improvement in the yield of aromatic compounds (25.5 wt% in test B2 compared to 17.8% in test T2 (the best result in the prior art)), with the highest absolute increase in p-xylene.

[0138] In C8 aromatic compounds, p-xylene is the dominant compound when ZSM-5 zeolite is dominant in the catalytic system, while in the case of Y zeolite (experimental T2), m-xylene is the dominant substance, possibly due to its smaller pore size than ZSM-5.

[0139] Table V

[0140] Operating conditions in CC3, yield (wt%), concentration and octane number of olefins, cycloalkane-olefins and / or dienes, aromatic compounds in naphtha produced in the reaction with biodiesel.

[0141]

[0142] Figures 5 to 10 The yields obtained in CC3 as a function of reaction temperature are shown. The optimal reaction temperature for maximizing the yield is between 460°C and 500°C. Lower reaction temperatures result in significantly increased yields of LCO, decanted oil, and coke, demonstrating the difficulty in loading vaporization.

[0143] This series of experiments also allows for the evaluation of the efficiency of coke combustion under extreme conditions with very low dense phase temperatures. Figure 11 The CO2 / CO ratio and carbon (RCC) in the regenerated catalyst are shown as a function of regeneration temperature (TFD). Under near-normal conditions, CC3 operates in a fully combusted manner; however, a significant trend toward partial combustion is observed as the dense phase temperature decreases.

[0144] Regarding RCC, catalyst regeneration remains sufficient even at lower temperatures, but it becomes quite high at 500°C, indicating some undercombustion under these conditions.

[0145] In demonstration-scale experiments A1 to A4, T1, and T2 for the production of bioaromatic compounds, ignition oil combustion was required to maintain the regeneration temperature at 600°C. However, experiments B1 to B6 demonstrated that both the regeneration and reaction temperatures could be reduced simultaneously, bringing the dense phase temperature down to 500°C, without affecting coke combustion on the catalyst. In this way, only the regenerator air preheater needs to be kept at 500°C to maintain thermal equilibrium. The combustion of coke formed on the catalyst is the reason for obtaining a temperature slightly higher than the 500°C provided by the air heater. For a coke yield of 0.8% by weight, calculations show an increase of approximately 80°C at the dense phase temperature. In this way, the dense phase temperature of the regenerator will be balanced between 550°C and 580°C, thus eliminating ignition oil combustion in the regenerator and therefore improving the economics of the process.

[0146] Therefore, a concentration and yield of aromatic compounds far exceeding those obtained using existing technologies were achieved: in this invention, the total aromatic compounds having 6 to 8 carbon atoms reached 25.5% by weight, and p-xylene reached 9.8% by weight, while for existing technologies, these values ​​only reached 13.5% and 2.3% by weight, respectively, and even then, they could not solve the problem of insufficient energy without the use of ignition oil combustion.

[0147] Example 3 Pilot-scale equipment using FCC CC3 tallow

[0148] 3.1- load:

[0149] Tallow was used as the loading and its properties are shown in Table VI.

[0150] Table VI

[0151] tallow characterization

[0152]

[0153] 3.2- catalyst :

[0154] The catalyst C5 based on ZSM-5 from Example 2 was used.

[0155] 3.3- Experimental setup:

[0156] The same CC3 pilot-scale apparatus as in Example 2 was used.

[0157] 3.4- analyze:

[0158] The same analysis and procedures as in Example 2 were used.

[0159] 3.5- Experimental results:

[0160] Table VII shows the yields of the products obtained from tallow using the present invention. The mass concentrations of olefins, cycloalkanes, and aromatic compounds in the produced naphtha were calculated using the PIANO method. High conversions exceeding 95% by weight were obtained in both experiments, D1 and D2. The yield of aromatic compounds with 6 to 8 carbon atoms in experiment D1 was higher than that obtained by any prior art experiment, reaching 28.2% by weight.

[0161] Example 4: FCC CC3 pilot plant using naphtha

[0162] 4.1- load:

[0163] Fossil-derived naphtha was used as the loading, and its properties are shown in Table VIII. Naphtha has a very low octane number, with MON and RON of only 46.5 and 53.8, respectively, making it unsuitable for use in Otto cycle engines.

[0164] 4.2- catalyst:

[0165] The catalyst C5 based on ZSM-5 from Example 2 was used.

[0166] 4.3- Experimental setup:

[0167] The same CC3 pilot-scale apparatus as in Example 2 was used.

[0168] 4.4- analyze:

[0169] The same analysis and procedures as in Example 2 were used.

[0170] The operating conditions in Table VII CC3, the yield (wt%), concentration, and octane number of olefins and aromatic compounds in the naphtha produced in the reaction with tallow.

[0171]

[0172] Table VIII

[0173] Naphtha characterization

[0174]

[0175] 4.5- Experimental results:

[0176] Table IX shows the yields of the products obtained from naphtha using this invention. The mass concentrations of olefins, cycloalkanes, and aromatic compounds in the produced naphtha were calculated using the PIANO method. In Experiment E1, the yield of aromatic compounds having 6 to 8 carbon atoms reached 16.6% by weight. Although this is lower than the value obtained using the renewable stream, it is significantly higher than the 9.4% by weight percentage associated with the naphtha fed into the process. Furthermore, the octane number of the produced naphtha is significantly higher than that of the naphtha fed into the process, approaching the value required for gasoline specifications.

[0177] Table IX

[0178] Operating conditions in CC3 using fossil naphtha as feedstock, yield (wt%), concentration, and octane number of olefins and aromatic compounds produced in the naphtha reaction.

[0179]

[0180] This invention lays the foundation for the use of low-value-added raw materials such as fatty acids and animal fats to produce aromatic petrochemical products from renewable sources in a profitable manner.

[0181] In addition to its excellent yield, this technology also enables the use of catalytic systems that are more sensitive to hydrothermal deactivation but more efficient for the production of light compounds and aromatic olefins, as it avoids high catalyst regeneration temperatures.

[0182] It should be noted that although the invention has been described with reference to the accompanying drawings, those skilled in the art can make modifications and adjustments as appropriate, provided that they are within the scope of the invention as defined herein.

Claims

1. A method for obtaining aromatic compounds from a stream comprising a straight chain having 5 to 18 carbon atoms, characterized in that... The method includes the following steps: a) A preheated feed stream (2) is introduced into a fluidized catalytic cracking reactor (5) and the preheated feed stream is contacted with a mesoporous zeolite catalyst (3) from a regenerator. The feed stream (2) is cracked at a reaction temperature of 440°C to 520°C to obtain a mixture of gaseous cracking products and coke-deactivated catalyst (6) from the fluidized catalytic cracking reactor (5), wherein the cracking products contain aromatic compounds and the catalyst-oil ratio is in the range of 5 to 30. b) Separate the cracking products and the deactivated catalyst from the coke from the mixture (6); c) The catalyst (6) that deactivates the coke separated in the previous step undergoes a distillation step (10) using steam (9) to remove light hydrocarbons (8). d) subjecting the catalyst (11) that has been distilled in the previous step to a regeneration step and burning the coke deposited on the catalyst particles, wherein the combustion temperature does not exceed 600°C, to obtain catalyst particles with higher activity than the spent catalyst. e) Distribute the regenerated catalyst with restored activity back to the fluidized catalytic cracking reactor (5) to continue the catalytic cracking process; f) Distribute the hydrocarbon stream (7) obtained in the catalytic cracking reactor to a fractionation section to separate the hydrocarbons according to the boiling point of the produced hydrocarbons; The catalyst regeneration temperature is at least 40°C and at most 100°C higher than the reaction temperature, and The air (13) used in the combustion of the coke deposited on the catalyst is preheated by the furnace (16) to supplement the thermal balance, and is thus fed into the regeneration section at a temperature of 400°C to 550°C.

2. The method according to claim 1, characterized in that... The preheated feed stream (2) contains triglycerides of C9 to C18 carbon chain fatty acids from plant and / or animal biomass.

3. The method according to claim 1, characterized in that... The preheated feed stream (2) contains free carbon chain fatty acids from C9 to C18.

4. The method according to claim 1, characterized in that... The preheated feed stream (2) contains triglycerides of fatty acids at a concentration greater than 65% by mass.

5. The method according to claim 2, characterized in that... The triglycerides are selected from soybean oil, castor oil, cottonseed oil, palm oil, pine oil, sunflower seed oil, jatropha oil, algae oil, or tallow.

6. The method according to claim 1, characterized in that... The preheated feed stream (2) contains a single oil or a mixture of two or more oils in any ratio.

7. The method according to claim 1, characterized in that... The preheated feed stream (2) is a stream of fossil hydrocarbons selected from refining streams with boiling points in the range of 40°C to 300°C.

8. The method according to claim 1, characterized in that... The mesoporous zeolite catalyst (3) is selected from the following structural types: MFI, MEL, ZSM-8, ZSM-12, ZSM-21, ZSM-23, ZSM-35, ZSM-38, IMF and TUN and any combination thereof.

9. The method according to claim 1, characterized in that... The fluidized catalytic cracking reactor (5) operates at a pressure ranging from 200 kPa to 400 kPa and a contact time ranging from 1.0 seconds to 3.0 seconds.

10. The method according to claim 1, characterized in that... The hydrocarbon stream (7) obtained in the fluidized catalytic cracking reactor (5) is fuel gas, liquefied gas, naphtha and other hydrocarbons.

11. The method according to claim 1, characterized in that... The fluidized catalytic cracking reactor (5) is a riser type.

12. An aromatic stream obtained by the method defined in claim 1, characterized in that... The aromatic stream has a typical naphtha distillation range of 30°C to 220°C and contains 70% to 90% by mass of aromatic compounds with benzene rings, a mixture of alkane and olefin compounds.

13. The aromatic hydrocarbon stream according to claim 12, characterized in that... The aromatic stream is a 100% renewable source.

14. The aromatic hydrocarbon stream according to claim 13, characterized in that... The aromatic stream has a minimum RON octane number of 100 and a minimum MON octane number of 85.

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