WO2024078904A1 - Gas-solid co-current downflow fluidised bed catalytic cracking process with oriented feed injector - Google Patents

Abstract

The invention relates to a gas-solid co-current downflow fluidised bed catalytic cracking process using: a vertical pipeline (1) carrying a downflow (4) of catalyst particles; a mixing chamber (2) comprising at least one first injector (5) of hydrocarbon feed (6); a gas-solid co-current downflow fluidised bed reactor (3) supplied with a catalyst and a hydrocarbon feed, the first injector having a predetermined angle α with respect to the horizontal so that: the vector of the resulting momentum (Mtot), calculated from the vector sum of the feed momentum (M6) and the catalyst momentum (M4), has an angle λ with respect to a horizontal plane of 70°-80°; and the ratio of the vertical component of the feed momentum (M6v) and the vertical and downflow momentum of the catalyst is between -0.2 and 0.1.

Description

CATALYTIC CRACKING PROCESS IN FLUIDIZED BED WITH DESCENDING GAS-SOLID CO-CURRENT WITH ORIENTED CHARGE INJECTOR

Technical area

The invention relates to the field of refining and petrochemicals and to processes and units for the chemical transformation of petroleum products, in particular hydrocarbon cuts, by Fluidized Bed Catalytic Cracking (“Fluid Catalytic Cracking” or FCC according to Anglo-Saxon terminology) for the production of light olefins (i.e., olefins comprising between 2 and 4 carbon atoms), and more particularly ethylene and propylene, and also aromatics (e.g. BTX), and more particularly paraxylene.

Prior art

The invention is part of the improvement of the design of the established flow zone of fluidized bed reactors with descending gas-solid co-current ("downer" or "down flow reactor" according to the Anglo-Saxon terminology), called hereinafter downflow reactors, used for example for high severity catalytic cracking (HS-FCC).

Ethylene, Propylene, Butene, Butadiene and aromatics such as Benzene, Toluene and Xylene (BTX) represent the basic products for the petrochemical industry. These products are generally obtained by catalytic reforming and/or thermal cracking (steam cracking) of hydrocarbons such as naphtha, kerosene or gas oil. These compounds are also obtained by fluidized bed catalytic cracking (FCC) of hydrocarbons, such as a Vacuum Distillate (DSV, or “vacuum gas oil” or VGO according to Anglo-Saxon terminology) and/or a residue ( under vacuum or atmospheric) distillation of hydrocarbons and/or naphtha, gas oils, complete crudes.

The high severity catalytic cracking (HS-FCC) process aims to increase yields of propylene and ethylene through high temperature reaction conditions, very short contact times (e.g. ~1 s), high flow rate ratios mass C of catalyst and mass flow rate O of charge (C/O).

The disadvantages associated with a conventional fluidized bed FCC reactor with ascending gas-solid co-current (“riser” according to Anglo-Saxon terminology) such as back-mixing (“back-mixing” according to Anglo-Saxon terminology) and the accumulation of catalyst in the vicinity of the wall with the consequence of overcracking of the hydrocarbons and the excessive formation of coke, hydrogen, methane and ethane do not make it possible to promote the production of olefins under high severity conditions.

In order to overcome these drawbacks, the HS-FCC process uses a downflow reactor, where the catalyst and feed are set in motion under gravity with a flow which approaches that of a piston type flow. The downward gas-solid flow in a reactor avoids back-mixing and over-cracking of products while the use of high C/O ratios ensures the predominance of catalytic reactions. The high temperature favors the formation of reaction intermediates such as light olefins while a controlled and short contact time avoids secondary reactions which are responsible for the consumption of such intermediates.

In contrast, downward gas-solid flow presents several major technological challenges, of which a significant challenge corresponds to the flow in the mixing chamber. Indeed, the initial mixture between catalyst and feed conditions the vaporization of the hydrocarbons and the gas-solid contact throughout the downflow reactor. The initial mixing is generally carried out in a fraction of a second with, for a typical HS-FCC, a flow rate of around 400 to 700 t/h of feed and 7000 to 21000 t/h of catalyst which requires efficient technology to have a mixing chamber that approximates a perfectly agitated zone.

Patent FR 2 753 453 B1 describes a downward flow cracking reactor comprising a contact zone between the hydrocarbons and the catalyst in which the injectors are oriented so as to direct droplets of charge against the current of the downward flow of catalyst particles, at an angle relative to the horizontal equal for example to 15°, but which can be between 2° and 45°.

Summary of the invention

In the context previously described, a first object of the present invention is to overcome the problems of the prior art and to provide a device for catalytic cracking in a fluidized bed with a descending gas-solid co-current with homogeneous catalyst flow, i.e. , in which the solid concentration in the cross section of the reactor is substantially uniform. Indeed, the device according to the invention makes it possible to obtain a homogeneous catalyst concentration between the central zone and the annular zone (i.e., in the proximity of the wall) of the descending gas-solid co-current reactor.

A second object of the present invention is to provide a device for catalytic cracking in fluidized bed with downward gas-solid co-current in which the concentration of the catalyst in the central zone is higher with less accumulation near the walls in the mixing chamber, which results in better contact with the vaporized load.

A third object of the present invention is to provide a device for catalytic cracking in a fluidized bed with a descending gas-solid co-current in which the arrangement of the injectors is optimized according to the C/O ratio. According to a first aspect, the aforementioned objects, as well as other advantages, are obtained by a device for catalytic cracking in a fluidized bed with a descending gas-solid co-current comprising, from top to bottom:

- a pipe adapted to transport a downward flow of catalyst particles;

- a mixing chamber connected to the pipe and adapted to be supplied by the pipe in a downward flow, the mixing chamber comprising at least one first hydrocarbon feed injector; And

- a fluidized bed reactor with descending gas-solid co-current connected to the mixing chamber and adapted to be supplied by the mixing chamber with a mixture comprising catalyst particles and hydrocarbon feed, in which the at least a first injector has a predetermined angle a relative to the horizontal meeting the following two criteria:

- the vector of the resulting momentum, calculated from the vector sum of the charge momentum and the catalyst momentum, has an angle À with respect to a horizontal plane of between 70° and 80° ; And

- the ratio of the vertical component of the quantity of load movement and the catalyst movement quantity is between -0.2 and 0.1.

Advantageously, the device makes it possible to homogenize the concentration of catalyst particles along the fluidized bed reactor. Advantageously, the device makes it possible to improve the contact between the catalyst and the vaporized charge. Advantageously, the device makes it possible to optimize the arrangement and orientation of the injectors according to the operating conditions envisaged.

According to one or more embodiments, the mixing chamber comprises between 2 and 12 first injectors, preferably between 3 and 8 first injectors.

According to one or more embodiments, the vector of the quantity of the resulting movement has an angle À with respect to a horizontal plane of between 72° and 77°; and/or the ratio of the vertical component of the vector of the momentum of the charge movement and the momentum of the catalyst is between -0.13 and 0.04.

According to one or more embodiments, the first injectors are arranged countercurrent to the descending flow at an angle a of between 15° and 45° relative to the horizontal. According to one or more embodiments, at least one first injector has an orientation offset relative to the diameter of the mixing chamber at an angle 0 greater than 0° and less than or equal to 45°, and preferably between 10° and 20°.

According to one or more embodiments, first injectors are arranged in one or more horizontal rows.

According to one or more embodiments, the mixing chamber comprises at least a second diluent injector.

According to one or more embodiments, the second injectors have an angle p relative to the horizontal of between 0° and 80°, and preferably between 10° and 45°.

According to one or more embodiments, second injectors are arranged in one or more horizontal rows.

According to one or more embodiments, the radial position of second injectors is in a separation space between the adjacent radial positions of two first injectors.

According to one or more embodiments, the radial position of at least a second injector is arranged relative to the radial position of an adjacent first injector at an angle 5 substantially equal to half of a separation angle y between two first adjacent injectors.

According to one or more embodiments, at least one second injector has an orientation offset relative to the diameter of the mixing chamber at an angle o greater than 0° and less than or equal to 45°, and preferably between 10° and 20°.

According to one or more embodiments, at least part of the mixing chamber comprises a central bulky part arranged substantially along a central/vertical axis of the mixing chamber and defining an annular orifice of the mixing chamber , whereby the catalyst particles spill and/or flow into the mixing chamber.

According to a second aspect, the aforementioned objects, as well as other advantages, are obtained by a process for catalytic cracking in a fluidized bed with a descending gas-solid co-current comprising the following steps:

- transport a downward flow of catalyst particles in a (vertical) pipe;

- supply a mixing chamber via the (vertical) pipe with the downward flow, the mixing chamber comprising at least one first hydrocarbon feed injector; - supply a fluidized bed reactor with a gas-solid co-current descending through the mixing chamber with a mixture comprising catalyst particles and hydrocarbon feed; And

- at least partially crack the hydrocarbon feed in the presence of the catalyst particles in the fluidized bed reactor with descending gas-solid co-current, to produce an effluent comprising at least partially coked catalyst and gaseous cracking products, in which the 'at least one first injector has a predetermined angle a relative to the horizontal meeting the following two criteria:

- the vector of the quantity of the resulting movement, calculated from the vector sum of the quantity of movement of charge and the quantity of movement (vertical and downward) of catalyst, has an angle À with respect to a horizontal plane between 70° and 80°; And

- the ratio of the vertical component of the quantity of load movement and the quantity of movement (vertical and downward) of catalyst, is between -0.2 and 0.1.

According to one or more embodiments, the method comprises at least one of the following operating conditions:

- The catalyst particles comprise a matrix made of clay, silica or silica alumina, optionally binder, optionally dopant, and/or zeolite, for example from 15% to 70% by weight of zeolite relative to the weight of the catalyst, preferably a Y zeolite and/or a ZSM-5 zeolite, very preferably a ZSM-5 zeolite, optionally doped;

- the grain density of the catalyst particles is between 1000 kg/m ³ and 2000 kg/m ³ , preferably between 1250 kg/m ³ and 1750 kg/m ³ .

- the hydrocarbon feedstock is a hydrocarbon feedstock with a start boiling temperature greater than or equal to 340°C or a hydrocarbon feedstock with an end boiling temperature less than or equal to 450°C;

- the downward flow of the catalyst particles in the (vertical) pipe upstream of the mixing chamber is in a dense fluidized regime and preferably with a mass flow greater than 200 kg/m ² s;

- reactor outlet temperature between 520°C and 750°C and preferably less than 650°C; absolute total pressure between 0.1 MPa and 0.5 MPa; - mass ratio of the catalyst to the C/O hydrocarbon load between 5° (kg/h)/(kg/h) and 35 (kg/h)/(kg/h) and preferably between 15 (kg/h) /(kg/h) and 30 (kg/h)/(kg/h);

- contact time te between the hydrocarbon feed and the catalyst less than 10 seconds, preferably between 0.5 seconds and 4 seconds;

- mass flow of catalyst particles of between 50 and 850 kg/(m ² s), preferably between 400 and 750 kg/(m ² s);

- gas surface speed between 2 m/s and 26 m/s, preferably between 6 m/s and 16 m/s;

- injection of diluent by the first injector and/or at least a second injector in a quantity representing 0% or 0.1% to 40% by weight, preferably 1% to 35% by weight and preferably between 1 % and 30% weight relative to the mass of the hydrocarbon filler.

Other characteristics and advantages of the invention of the aforementioned aspects will appear on reading the description below and non-limiting examples of embodiments, with reference to the figures appended and described below.

List of Figures

Figure 1 represents a section diagram of an FCC device according to one or more embodiments of the present invention comprising injectors for the homogenization of the catalyst flow.

Figure 2 shows a schematic and sectional view showing fluid flow in the mixing chamber of an FCC device according to one or more embodiments of the present invention.

Figure 3 shows a schematic top view of the mixing chamber of an FCC device according to one or more embodiments of the present invention.

Figure 4 shows 3D views of an FCC device according to one or more embodiments of the present invention.

Figure 5 shows 3D views of a diagram A of a reference FCC device, and of a diagram B of an FCC device according to one or more embodiments of the present invention.

Figure 6 shows 3D views of diagrams A and B of the volume fraction of the catalyst averaged over time in the reference FCC device of Figure 5, and in the FCC device according to the invention of Figure 5, respectively. Figure 7 shows the radial profiles A and B of the fraction of the solid and mass flow of the solid at 2.8 m below the injectors of the reference FCC device of Figure 5, and of the FCC device according to the invention of the Figure 5, respectively.

Detailed description of the invention

Embodiments according to the aforementioned aspects will now be described in detail. In the following detailed description, numerous specific details are set forth in order to provide a more in-depth understanding of the device and method according to the present invention. However, it will appear to those skilled in the art that the device can be implemented without these specific details. In other cases, well-known features have not been described in detail to avoid unnecessarily complicating the description.

In this description, the term “comprising” is synonymous with (means the same as) “comprising”, “include” and “contain”, and is inclusive or open and does not exclude other elements not recited. It is understood that the term “understand” includes the exclusive and closed term “consist”. Furthermore, in the present description, the terms "essentially" or "substantially" or "approximately" correspond to an approximation of ± 10%, preferably ± 5%, most preferably ± 1%.

The invention relates to a device and a method for fluidized bed catalytic cracking for the chemical transformation of petroleum products (FCC), used for example for high severity catalytic cracking (HS-FCC).

An FCC unit generally treats a hydrocarbon cut (called heavy) from the vacuum distillation unit as a vacuum gas oil or a vacuum residue, or even an atmospheric residue, alone or in a mixture. An FCC unit can also process a hydrocarbon cut (called light) such as a gasoline cut or a diesel cut, alone or in a mixture. It is also possible to process a mixture of light and heavy hydrocarbon cuts, or even a complete crude. In order to increase the yields of propylene and ethylene using high severity reaction conditions (high temperature, very short contact times, high C/O ratios between the catalyst flow rate C and the feed flow rate O), the devices and catalytic cracking processes generally use a fluidized bed reactor with a descending gas-solid co-current (“downer” or “down flow reactor” according to Anglo-Saxon terminology), hereinafter called a downward flow reactor.

However, descending gas-solid flow sometimes presents several technological challenges, including achieving homogeneous flow and mixing in the steady flow section of the descending reactor. In order to overcome these drawbacks, improvements in the downflow reactor technology described in patent FR 2 753 453 B1 have been identified, to respond to the challenge presented above through a specific arrangement of hydrocarbon feed injectors, and optionally diluent injectors, to improve contact between the catalyst and the hydrocarbons in the mixing chamber.

The device according to the invention

With reference to Figure 1, the device according to one or more embodiments of the present invention comprises from top to bottom:

- a pipe 1 (substantially vertical);

- a mixing chamber 2; And

- a downward flow reactor 3 (substantially vertical).

Line 1 is suitable for supplying mixing chamber 2 with solid catalyst (particles). Line 1 mainly transports solids, as well as a fluidization gas entrained by the descending solid. Pipe 1 presents a flow like a supply column (“standpipe” according to Anglo-Saxon terminology) well known to those skilled in the art.

The mixing chamber 2 is connected to pipe 1 and is adapted to feed the downflow reactor 3 with a mixture comprising catalyst particles, a hydrocarbon feedstock and optionally a diluent.

The downflow reactor 3 is connected to the mixing chamber 2 and is adapted to at least partially crack the hydrocarbon feed in the presence of the catalyst particles to produce an effluent comprising at least partially coked catalyst and gaseous cracking products, and optionally unconverted vaporized charge.

Specifically, with reference to Figure 1, the pipe 1 supplies the mixing chamber 2 with a descending flow 4 of (hot) catalyst particles, the mixing chamber 2 comprising one or more first injectors 5 of hydrocarbon feed 6. According to one or more embodiments, the first injector(s) 5 is adapted to inject diluent (eg water vapor) with the load. According to one or more embodiments, the mixing chamber 2 comprises one or more second injectors 7 of diluent 8. In the mixing chamber 2, the descending flow 4 comes into contact with the hydrocarbon feed 6 atomized using the or first injectors 5 and optionally with diluent 8, introduced for example by the second injector(s) 7. Advantageously, the injection of diluent 8 makes it possible to reduce the partial pressure of the hydrocarbon feed and to reduce secondary reactions. Advantageously, the injection of diluent 8 makes it possible to improve the atomization of the hydrocarbon feedstock 6. According to one or more embodiments, the diluent 8 is chosen from the group consisting of water vapor, nitrogen , CO2, light hydrocarbons (eg C1-C5 compounds), combustion fumes. According to one or more embodiments, the diluent 8 comprises or consists of water vapor.

According to one or more embodiments, the pipe 1 has a constant section geometry, such as cylindrical, square, rectangular or hexagonal, or a variable section, such as a truncated pyramid or cone, or a combination of different geometric shapes. According to one or more embodiments, the pipe 1 is cylindrical in shape and optionally of variable diameter. According to one or more embodiments, the pipe 1 is at least partially frustoconical in shape. According to one or more embodiments, the pipe 1 comprises (in the direction of the flow of the solid catalyst): a cylindrical section, for example whose diameter is chosen to obtain a solid flow of 100 to 800 kg/(m ² s) and preferably between 300 to 600 kg/(m ² s); a frustoconical section (called a narrowing section) adjacent to the mixing chamber 2, the diameter of which decreases, for example in order to achieve a solid flow, without taking into account the internals, between 400 and 2000 kg/(m ² s) and preferably between 700 and 1500 kg/(m ² s); and optionally a second cylindrical section, for example whose diameter is chosen to obtain a solid flow, without taking into account the internals, between 400 and 2000 kg/(m ² s) and preferably between 700 and 1500 kg/(m ² s).

According to one or more embodiments, the mixing chamber 2 has a constant section geometry, such as cylindrical, square, rectangular or hexagonal, or a variable section, such as a truncated pyramid or cone, or a combination of different geometric shapes. . The mixing chamber can be equipped with central internals, such as a central bulk part described below with reference to Figure 5. According to one or more embodiments, the mixing chamber 2 is of cylindrical shape and optionally of variable diameter. According to one or more embodiments, the mixing chamber 2 is at least partially frustoconical in shape. According to one or more embodiments, the mixing chamber 2 comprises an upper limit of section S1 connecting the mixing chamber 2 to pipe 1 and a lower limit of section S2 connecting the mixing chamber 2 to the downflow reactor 3, the S1/S2 ratio being less than 0.9 and preferably less than 0.7. According to one or more embodiments, the S1/S2 ratio is between 0.4 and 0.9, preferably between 0.5 and 0.7.

According to one or more embodiments, the mixing chamber 2 comprises between 2 and 12 first injectors 5, preferably between 3 and 8 first injectors 5. With reference to Figures 1 and 2, the applicant has identified that the axial direction of the first injectors 5 is a key element for properly dispersing the catalyst. Specifically according to the invention, the first injectors 5 have a predetermined angle a (alpha) relative to the horizontal meeting the following two criteria:

- the vector of the quantity of the resulting movement M _tot , calculated from the vector sum of the quantity of movement (“momentum” according to Anglo-Saxon terminology; defined as mass (or mass flow) * speed) of charge M ₆ and the momentum of catalyst M ₄ , has an angle À (lambda) relative to the horizontal between 70° and 80° and preferably between 72° and 77°; And

- the ratio of the vertical component of the vector of the quantity of movement of charge Mev and the quantity of movement of catalyst M ₄ , is between -0.2 and 0.1 and preferably between -0.13 and 0.04 .

Compliance with the two criteria mainly gives axial directions of the first injectors 5 in counter-current of the descending flow 4, for example at an angle a between 0° and 80°, or slightly in co-current of the descending flow 4, for example at an angle between -20° and 0°. According to one or more embodiments, the first injectors 5 are arranged against the current of the descending flow 4 at an angle a of between 15° and 45° relative to the horizontal.

Furthermore, with reference to Figure 3, to promote turbulence in the mixing chamber 2 and direct the charge towards areas with a higher concentration of catalyst, at least a first injector 5 can have an orientation (according to a projection on a horizontal plane) offset from the diameter of the mixing chamber 2 at an angle 0 (theta) greater than 0° and less than or equal to 45°, and preferably between 10° and 20°. According to one or more embodiments, first adjacent off-axis injectors 5 have respectively positive and negative angles 0, in particular to improve turbulence.

With reference to Figure 4, according to one or more embodiments, first injectors 5 are arranged in one or more horizontal rows, i.e., perpendicular to the central/vertical axis Z of the pipe 1, of the mixing chamber 2 and the downflow reactor 3. According to one or more embodiments, the first adjacent injectors 5 of a row of first injectors 5 are arranged on the wall of the mixing chamber 2 at a separation angle y (gamma) being substantially equal to 360° divided by the number of first injectors 5 of said row.

According to one or more embodiments, the mixing chamber 2 comprises between 2 and 12 second injectors 7, preferably between 3 and 8 second injectors 7. With reference to Figure 1, the second injectors 7 have an angle p (beta) relative to the horizontal of between 0° and 80°, and preferably between 10° and 45°.

With reference to Figure 4, according to one or more embodiments, second injectors 7 are arranged in one or more horizontal rows (perpendicular to the central/vertical axis Z).

With reference to Figure 4, according to one or more embodiments, second injectors 7 are arranged below (e.g. in a row) first injectors 5. According to one or more embodiments, second injectors 7 are arranged between two rows of first injectors 5.

With reference to Figures 3 and 4, according to one or more embodiments, the radial position of second injectors 7 is in a separation space between the adjacent radial positions of two first injectors 5. According to one or more embodiments, the position radial of at least a second injector 7 is arranged relative to the radial position of a first injector 5 adjacent a second injector 7 (according to a projection on a horizontal plane) at an angle 5 (delta) substantially equal to half of the separation angle y between two first adjacent injectors 5.

With reference to Figures 3 and 4, to promote turbulence in the mixing chamber 2 and direct the catalyst towards areas with higher charge concentration, the second injectors 7 can have an orientation (according to a projection on a horizontal plane) offset relative to the diameter of the mixing chamber at an angle o (sigma) greater than 0° and less than or equal to 45°, and preferably between 10° and 20°. To make the top view in Figure 4 clearer, a row of injectors 5 has been removed.

With reference to Figure 1, the mixing chamber 2 feeds the downflow reactor 3 with a mixture of hydrocarbon feed 6, catalyst particles and optionally diluent 8. Advantageously, the hydrocarbon feed 6 and the catalyst particles give rise to to the cracking reactions which are completed in the downward flow reactor 3 of a height L (along the central/vertical axis Z) to produce a hydrocarbon effluent 9 comprising cracking products, spent catalyst and potentially part of the unreacted hydrocarbon load.

With reference to Figure 5, according to one or more embodiments, at least part of the mixing chamber 2 comprises a central bulky part 10 ("plug" according to Anglo-Saxon terminology) arranged substantially along the central/vertical axis Z and defining an annular orifice 11 of the mixing chamber 2, through which the catalyst particles spill and/or flow into the mixing chamber 2. According to one or more embodiments, the vertical position of the first injectors 5 and/or the second injectors 7 is between the upper end and the lower end of the central bulk part 10, ie, the load 6 and optionally the diluent 8 are introduced into the annular orifice 11 of the mixing chamber 2.

According to one or more embodiments, the downward flow reactor 3 has a constant section geometry, such as cylindrical, square, rectangular or hexagonal, preferably cylindrical. According to one or more embodiments, the downward flow reactor 3 is cylindrical in shape and optionally of variable diameter. According to one or more embodiments, the diameter of the downward flow reactor 3 is defined such that the superficial speed of the gas passing through it is between 2 m/s and 26 m/s, preferably between 6 m/s and 16 m/s. s.

The catalyst

The catalyst is a solid catalyst (e.g. particles of density, size and grain shape chosen for use in a fluidized bed). The densities, sizes and shapes of the catalysts for fluidized beds are known to those skilled in the art, and will not be described further. The catalyst may be any type of catalytic cracking catalyst.

According to one or more embodiments, the catalyst is an FCC type catalyst, containing for example what is commonly called a matrix made of clay, silica or silica alumina, optionally binder, and/or zeolite, for example example from 15% to 70% by weight of zeolite relative to the weight of the catalyst, preferably a Y zeolite and/or a ZSM-5 zeolite. According to one or more embodiments, the catalyst comprises a ZSM-5 zeolite. According to one or more embodiments, the grain density of the catalyst is between 1000 kg/m ³ and 2000 kg/m ³ . According to one or more embodiments, the grain density of the catalyst is between 1250 kg/m ³ and 1750 kg/m ³ .

According to one or more embodiments, the catalyst comprises at least one binder (e.g. from 30% to 85% by weight) chosen from alumina, silica, silica-alumina, magnesia, titanium oxide, zirconia , clays and boron oxide, alone or in a mixture and preferably among silica, silica-alumina and clays, alone or in a mixture.

According to one or more embodiments, the catalyst comprises at least one doping element (e.g. from 0 to 10% by weight) chosen from phosphorus, magnesium, sodium, potassium, calcium, iron, boron, manganese , lanthanum, cerium, titanium, tungsten, molybdenum, copper, zirconium and gallium, alone or in mixture.

According to one or more embodiments, the catalyst comprises and/or consists of zeolite, such as ZSM-5, optionally doped. Load

According to one or more embodiments, the hydrocarbon feed 6 is a hydrocarbon feed (so-called heavy), characterized by a start-of-boiling temperature of substantially 340°C, or even greater than 340°C, often greater than 380°C, such as a heavy hydrocarbon cut, for example from a vacuum distillation unit, such as vacuum gas oil/distillate (“vacuum gas oil” or “VGO” according to Anglo-Saxon terminology) or a vacuum residue , an atmospheric residue, a vacuum gas oil from a conversion unit, such as a coking gas oil ("Heavy Coker Gas Oil" or "HCGO" according to Anglo-Saxon terminology) or a heavy hydrocarbon cut from a bubbling bed or entrained bed hydroconversion unit (such as the H-Oil, LC-Fining, EST, VCC or Uniflex processes), a recycle of a hydrocracking stage, alone or in a mixture.

According to one or more embodiments, the hydrocarbon feedstock 6 is a so-called light feedstock, characterized by an end of boiling temperature less than or equal to 450°C, often less than 400°C, such as a gasoline cut or a gas oil cut, for example from an atmospheric distillation unit, or from a conversion unit, such as gasoline or gas oil from a hydrocracking unit, or gasoline or gas oil from a processing unit. coking or a gasoline or gas oil from a bubbling bed or entrained bed hydroconversion unit (such as the H-Oil, LC-Fining, EST, VCC or Uniflex processes), or a gasoline or gas oil from an FCC unit, or a recycle of the FCC unit in question, alone or in a mixture. According to one or more embodiments, it is also possible to process a mixture of light and heavy hydrocarbon cuts, or even a complete crude.

On contact with the descending flow 4 of hot catalyst particles, the pulverized hydrocarbon feed 6 vaporizes and endothermic cracking reactions occur along the descending flow reactor 3, thus reducing the temperature and producing:

- recoverable products (e.g. C1-C4 gas including olefins; a gasoline cut including aromatics);

- optionally a light diesel cut (“Light Cycle Oil” or LCO according to Anglo-Saxon terminology) ⁰ ;

- optionally a heavy diesel cut (“Heavy Cycle Oil” or HCO according to Anglo-Saxon terminology);

- optionally an oil in the form of mud (“slurry” according to Anglo-Saxon terminology); and optionally a solid residue (coke) adsorbed on the catalyst. The process according to the invention

The process according to the present invention comprises a catalytic cracking step for the production of light olefins (and in particular ethylene and propylene), aromatics (and in particular benzene, toluene and xylenes), and gasoline (and optionally LCO, HCO and slurry), by catalytic cracking of the hydrocarbon feed 6 (fed by the first injector 5) by contact with the descending flow 4 of hot catalyst particles (fed by line 1), and optionally the diluent 8 (supplied by the second injector 7) in the mixing chamber 2 then in the downflow reactor 3.

According to one or more embodiments, the downward flow 4 of the catalyst particles in line 1 upstream of the mixing chamber 2 is in a dense fluidized regime and preferably with a mass flow greater than 200 kg/m ² s, for for example preferably allowing a regime with descending bubbles.

In the present application, the term “dense fluidized bed” means a gas-solid fluidized bed operating in a homogeneous regime, in a bubbling regime or in a turbulent regime.

In the present application, the term “homogeneous fluidized bed” means a gas-solid fluidized bed whose gas speed is between the minimum fluidization speed and the minimum bubbling speed. These speeds depend on the properties of the solid catalyst (density, size, shape of the grains, etc.). The solid volume fraction is between a value close to 0.45 and the maximum solid volume fraction corresponding to a fixed, non-fluidized bed, generally close to 0.6.

In the present application, the term “bubbling fluidized bed” means a gas-solid fluidized bed whose gas speed is between the minimum bubbling speed and the transition speed to the turbulent regime. These speeds depend on the properties of the solid catalyst (density, size, shape of the grains, etc.). The volume fraction of solid is between a value close to 0.35 and a value close to 0.45.

In the present application, the term “turbulent fluidized bed” means a gas-solid fluidized bed whose gas speed is between the transition speed to the turbulent regime and the transport speed. The volume fraction of solid is between a value close to 0.25 and a value close to 0.35.

In the present application, the term “transported fluidized bed” means a gas-solid fluidized bed whose gas velocity is greater than the transport velocity. The solid volume fraction is less than a value close to 0.25. In the present application, the term “speed of transport” corresponds to the speed with which essentially all the solid is carried away by the gas.

According to one or more embodiments, the injectors 5 are adapted to atomize the hydrocarbon feed 6 (liquid) and penetrate the catalyst flow.

According to one or more embodiments, the operating conditions of the pipe 1 and/or the downflow reactor 3 are chosen from the following conditions:

- temperature (reactor outlet) between 520°C and 750°C and preferably less than 650°C;

- absolute total pressure between 0.1 MPa and 0.5 MPa;

- mass ratio of catalyst 4 to hydrocarbon feedstock 6 C/O between 5 (kg/h)/(kg/h) and 35 (kg/h)/(kg/h) and preferably between 15 (kg/h) )/(kg/h) and 30 (kg/h)/(kg/h);

- contact time t _c between the hydrocarbon feed 6 and the catalyst less than 10 seconds, preferably between 0.5 seconds and 4 seconds;

- mass flow of catalyst particles of between 50 and 850 kg/(m ² s), preferably between 400 and 750 kg/(m ² s); And

- gas surface speed between 2 m/s and 26 m/s, preferably between 6 m/s and 16 m/s.

In the present description, the contact time t _c is defined as the product of the solid volume fraction E _S by the bed height H _s (eg reactor height L), divided by the superficial gas velocity vsg, this integrating all along the height of the bed, as defined below in the mathematical formula Math 1.

Math 1

According to one or more embodiments, a quantity of diluent 8 (e.g. nitrogen and/or water vapor) is added to the charge to reduce the partial pressure of hydrocarbons of the charge and the diluent is introduced at a rate of one quantity representing 0% or 0.1% to 40% by weight, preferably 1% to 35% by weight and preferably between 1% and 30% by weight relative to the mass of the hydrocarbon filler 6.

According to one or more embodiments, at the end of the catalytic cracking step in the downflow reactor 3, the gaseous products and the catalyst, and optionally the feed unconverted vaporized, are separated in the gas/solid separator (not shown) enclosing a dense fluidized bed where the cracking reactions can continue.

According to one or more embodiments, the operating conditions of the separator are chosen from the following conditions:

- temperature (reactor outlet) between 500°C and 750°C, preferably between 550°C and 700°C, even more preferably between 580°C and 685°C;

- absolute total pressure between 0.1 MPa and 0.5 MPa and preferably between 0.1 MPa and 0.4 MPa and preferably between 0.1 MPa and 0.3 MPa;

- mass ratio of the catalyst to the feed (unconverted vaporized feed and gaseous products) C/O between 5 (kg/h)/(kg/h) and 40 (kg/h)/(kg/h);

- contact time t _c between the charge and the catalyst of between 500 milliseconds (ms) and 10 seconds; And

- partial pressure of the hydrocarbons of the charge (PPHcharge) between 0.01 MPa and 0.3 MPa, preferably between 0.02 MPa and 0.2 MPa and preferably between 0.05 MPa and 0.15 MPa.

According to one or more embodiments, at the outlet of the separator the coked catalyst is sent to an optional stripper (not shown) to strip the hydrocarbons remaining adsorbed on the surface of the catalyst by means of a second diluent.

According to one or more embodiments, the operating conditions of the stripper are chosen from the following conditions:

- residence time of the catalyst in the stripper: between 10 seconds and 180 seconds, preferably between 30 seconds and 120 seconds;

- superficial gas speed between the minimum fluidization speed and the transition speed to the turbulent regime, for example between 0.01 m/s and 0.5 m/s, preferably between 0.15 m/s and 0.4 m /s ;

- solid flow between 25 kg/m ² s and 200 kg/m ² s, preferably between 50 kg/m ² s and 150 kg/m ² s and preferably between 50 kg/m ² s and 100 kg/m ² s ;

- temperature between 500°C and 750°C, preferably between 550°C and 650°C;

- absolute total pressure between 0.1 MPa and 0.5 MPa and preferably between 0.1 MPa and 0.4 MPa and preferably between 0.1 MPa and 0.3 MPa;

- solid volume fraction between 0.25 and 0.6, preferably between 0.4 and 0.6. According to one or more embodiments, at the outlet of the separator or stripper, the coked solid is transported into a regenerator (not shown) in which an air supply burns the coke from the catalyst to produce a hot regenerated catalyst and combustion gases , the hot regenerated catalyst being able to supply the descending flow 4 with hot catalyst particles.

According to one or more embodiments, the operating conditions of the regenerator are chosen from the following conditions:

- superficial gas speed between 0.1 m/s and 2 m/s, preferably 0.2 m/s and 1.5 m/s;

- catalyst residence time between 30 seconds and 20 minutes, preferably between 1 minute and 10 minutes;

- temperature between 500°C and 840°C, preferably between 650°C and 750°C.

Examples

With reference to Figure 5, the flows in a reference device A and a device according to the invention B operating under reaction-free conditions were compared in order to study the hydrodynamics.

The reference device A includes:

- a pipe 1, composed of a cylindrical section, a narrowing cone and a cylindrical section;

- a frustoconical mixing chamber 2 (S1/S2 being less than 1) comprising a central bulk part 10 defining an annular orifice 11;

- a cylindrical downflow reactor 3 having a length of 4.07 m and an internal diameter of 0.42 m;

- four first injectors 5 of hydrocarbon feed 6 positioned in co-current of the downward flow 4 of catalyst particles with an angle of 45° downwards relative to the horizontal direction, which produces an angle À of the vector of the quantity of the resulting movement M _tot of 79° relative to the horizontal. The ratio of the vertical component of the charge momentum M _6v and the catalyst momentum M ₄ is 0.23 which is outside the recommended range; and four second injectors 7 of diluent 8 (water vapor) positioned in co-current of the downward flow 4 of catalyst particles with an angle of 45° downwards with respect to the horizontal direction. The device according to the invention B comprises:

- a pipe 1, composed of a cylindrical section, a narrowing cone and a cylindrical section;

- a frustoconical mixing chamber 2 (S1/S2 being less than 1) comprising a central bulk part 10 defining an annular orifice 11;

- a cylindrical downflow reactor 3 having a length of 4.07 m and an internal diameter of 0.42 m;

- four first injectors 5 of hydrocarbon feed 6 positioned counter-current to the downward flow 3 of catalyst particles with an angle a of 30° upwards relative to the horizontal direction, which produces an angle À of the quantity vector of the resulting movement M _tot 72° relative to the horizontal. The ratio of the vertical component of the charge momentum Mev and the catalyst momentum M4 is -0.16; And

- four second injectors 7 of diluent 8 (water vapor) positioned countercurrent to the downward flow 3 of catalyst particles with an angle p of 30° upwards relative to the horizontal direction.

The configurations of the reference device A and the device according to the invention B were simulated in CFD with the Barracuda© tool under the following operating conditions:

- the catalyst flow of 607 kg/m ² s;

- the catalyst has a diameter dso of 73 pm and a grain density of 1418 kg/m ³ (ie, group A of the Geldart classification);

- the air flow rate of 1.85 kg/s presents a ratio between the first injectors 5 to the second injectors 7 of 70/30;

- the first injectors 5 are positioned 0.2 m above the second injectors 7;

- the angle 5 of rotation between the first injectors 5 and the second injectors 7 is 45°;

- the gas speed leaving the injectors is 90 m/s for the first injectors 5 and 76 m/s for the second injectors 7;

- the flow is under ambient conditions without reaction;

- a central bulky part 10 is positioned in the center of the mixing chamber 2.

Figure 6 presents the Volume Fraction of Particles, denoted FVP, for the two configurations A and B of the reference device A and the device according to the invention B, respectively, over 13 sections. The first two sections are respectively at the height of the first and second injectors 5 and 7, and the following sections are spaced 0.35 m apart. Advantageously, the radial distribution of the solid throughout the downward flow reactor 3 is always more homogeneous for the device according to the invention B compared to the reference device A. In addition, it should be noted a significant concentration of the solid at the proximity of the wall for the reference device A in relation to the device according to the invention B. Figure 7 shows the radial profiles A and B of the volume fraction of the particles FVP, and of the Mass Flux of the Particles, denoted FM P, of the reference device A and the device according to the invention B of Figure 5, respectively. The radial profiles A and B are produced in a direction x (perpendicular to the central/vertical axis Z) at a height of 2.8 m below the first injectors 5. It can be noted that the device according to the invention B produces a more homogeneous radial profile B with better phase distribution. The coefficient of variation of the volume fraction values of the FVP particles on these radial profiles A and B is 40% for the reference device A and 15% for the device according to the invention B. Equivalently, the coefficient of variation of the mass flow values of the FMP particles for profiles A and B is 59% for the reference device A and 16% for the device according to the invention B, which indicates better dispersion of the catalyst on the section for the device according to the invention B.

Claims

Claims 1. Process for catalytic cracking in a fluidized bed with a descending gas-solid co-current comprising the following steps: - transport a downward flow (4) of catalyst particles in a vertical pipe (1); - supply a mixing chamber (2) via the vertical pipe (1) with the downward flow (4), the mixing chamber (2) comprising at least one first injector (5) of hydrocarbon feed (6); - supply a fluidized bed reactor with descending gas-solid co-current (3) through the mixing chamber (2) with a mixture comprising catalyst particles and hydrocarbon feed; And - at least partially crack the hydrocarbon feed (6) in the presence of the catalyst particles in the fluidized bed reactor with descending gas-solid co-current (3), to produce an effluent (9) comprising at least partially coked catalyst and gaseous cracked products, in which the at least one first injector (5) has a predetermined angle a relative to the horizontal meeting the following two criteria: - the vector of the resulting movement quantity (M _tot ), calculated from the vector sum of the load movement quantity (Me) and the vertical and downward catalyst movement quantity (M4), has an angle À relative to a horizontal plane between 70° and 80°; And - the ratio of the vertical component of the quantity of load movement (Mev) and the quantity of vertical and downward movement of catalyst (M ₄ ), is between -0.2 and 0.1. 2. Method according to claim 1, wherein the mixing chamber (2) comprises between 2 and 12 first injectors (5), preferably between 3 and 8 first injectors (5). 3. Method according to claim 1 or claim 2, in which the vector of the quantity of the resulting movement (M _tot ) has an angle À with respect to a horizontal plane of between 72° and 77°; and/or the ratio of the vertical component of the vector of the amount of charge movement (Mev) and the amount of vertical and downward movement of catalyst (M ₄ ), is between -0.13 and 0.04. 4. Method according to any one of the preceding claims, in which the first injectors (5) are arranged against the current of the descending flow (4) at an angle a between 15° and 45° relative to the horizontal. 5. Method according to any one of the preceding claims, in which at least one first injector (5) has an orientation offset relative to the diameter of the mixing chamber (2) at an angle 0 greater than 0° and less than or equal to at 45°, and preferably between 10° and 20°. 6. Method according to any one of the preceding claims, in which first injectors (5) are arranged in one or more horizontal rows. 7. Method according to any one of the preceding claims, wherein the mixing chamber (2) comprises at least one second injector (7) of diluent (8). 8. Method according to claim 7, in which the second injectors (7) are arranged countercurrent to the downward flow (4) and have an angle p relative to the horizontal of between 0° and 80°, and preferably between 10° and 45°. 9. Method according to claim 7 or claim 8, wherein second injectors (7) are arranged in one or more horizontal rows. 10. Method according to any one of claims 7 to 9, in which the radial position of second injectors (7) is in a separation space between the adjacent radial positions of two first injectors (5). 11. Method according to any one of claims 7 to 10, in which the radial position of at least one second injector (7) is arranged relative to the radial position of a first injector (5) adjacent at an angle 5 substantially equal to half of a separation angle y between two adjacent first injectors (5). 12. Method according to any one of claims 7 to 11, in which at least one second injector (5) has an orientation offset relative to the diameter of the mixing chamber (2) at an angle o greater than 0° and less or equal to 45°, and preferably between 10° and 20°. 13. Method according to any one of the preceding claims, in which at least part of the mixing chamber (2) comprises a central bulk part (10) arranged substantially along a central/vertical axis (Z) of the mixing chamber (2) and defining an annular orifice (11) of the mixing chamber (2), through which the catalyst particles discharge and/or flow into the mixing chamber (2). 14. Method according to claim 1, comprising at least one of the following operating conditions: - the catalyst particles comprise a matrix made of clay, silica or silica alumina, optionally binder, optionally dopant, and/or zeolite, for example from 15% to 70% by weight of zeolite relative to the weight of the catalyst, preferably a Y zeolite and/or a ZSM-5 zeolite, very preferably a ZSM-5 zeolite, optionally doped; - the grain density of the catalyst particles is between 1000 kg/m ³ and 2000 kg/m ³ , preferably between 1250 kg/m ³ and 1750 kg/m ³ ; - the hydrocarbon feedstock (6) is a hydrocarbon feedstock with a start boiling temperature greater than or equal to 340°C or a hydrocarbon feedstock with an end boiling temperature less than or equal to 450°C; - the downward flow (4) of the catalyst particles in the vertical pipe (1) upstream of the mixing chamber (2) is in a dense fluidized regime and preferably with a mass flow greater than 200 kg/m ² s; - reactor outlet temperature between 520°C and 750°C and preferably less than 650°C; - absolute total pressure between 0.1 MPa and 0.5 MPa; - mass ratio of the catalyst (4) to the hydrocarbon feed (6) C/O between 5 (kg/h)/(kg/h) and 35 (kg/h)/(kg/h) and preferably between 15 (kg/h)/(kg/h) and 30 (kg/h)/(kg /h); contact time te between the hydrocarbon feed (6) and the catalyst less than 10 seconds, preferably between 0.5 seconds and 4 seconds; - mass flow of catalyst particles of between 50 and 850 kg/(m ² s), preferably between 400 and 750 kg/(m ² s); - gas surface speed between 2 m/s and 26 m/s, preferably between 6 m/s and 16 m/s; - injection of diluent (8) by the first injector (5) and/or at least a second injector (7) in a quantity representing 0% or 0.1% to 40% by weight, preferably 1% to 35 % by weight and preferably between 1% and 30% by weight relative to the mass of the hydrocarbon filler (6).