Causes of deactivation and methods of regeneration of catalysts. Interaction of catalysts with the reaction medium

Changes in the composition of catalysts during the reaction can be as follows: 1 chemical changes leading to phase transformations of the active component; 2 changes in volumetric composition without phase transformations; 3 changes in the composition of the surface layer of the catalyst. The impact of the reaction medium can lead to a change in the ratio of the components included in the catalyst, as well as to the dissolution of new components or partial removal of old ones. The stable composition of the catalyst is determined by the ratio of the rates of binding or consumption...

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Interaction of catalysts with the reaction medium.

Causes of deactivation and methods of regeneration of catalysts.

6.1 Interaction of catalysts with the reaction medium.

The final properties of the catalysts are formed under the action of the reaction medium. Changes in the composition of catalysts during the reaction can be as follows:

1) chemical changes leading to phase transformations of the active ingredient;

2) changes in the volumetric composition without phase transformations;

3) changes in the composition of the surface layer of the catalyst.

The impact of the reaction medium can lead to a change in the ratio of the components that make up the catalyst, as well as to the dissolution of new components or partial removal of old ones.

The stable composition of the catalyst is determined by the ratio of the rates of binding or consumption of a certain component of the catalyst as a result of interaction with the reactants. In accordance with the change in the degree of conversion of the reactants, the stationary composition of the catalyst, and, consequently, its properties can change significantly along the catalyst bed in the reactor.

The time to reach the steady state can be very long under appropriate conditions, such as low temperature. The rate of establishment of the stationary composition of the catalyst in a certain temperature range depends on whether the activity measurements are approaching the temperature from the side of higher or from the side of lower temperature.

Substances that are catalysts are poisoned by poisons or inhibitors. Poisons completely reduce the activity of the catalyst, while inhibitors partially suppress the activity and can change the selectivity of the catalysts. Solid catalysts on the outer and inner (inside the pores) surface have active centers - polyhedra. These active sites on the surface are distributed by energy (activated adsorption and chemisorption), and their share of the total catalyst surface ranges from 1 to 10 percent. On this basis, complete poisoning of the catalyst, i.e., reducing their catalytic activity to zero, also requires a small amount of poison. Consequently, the poison molecules, when adsorbed on the surface of the catalyst, do not cover the entire surface of the catalyst, but only its active part, active centers, thereby poisoning the effect of the entire catalyst for the chemical process.

6.2 Reasons for decontamination. Catalyst poisoning.

Catalyst poisoning is the partial or complete loss of activity due to the action of a small amount of substances called contact poisons or inhibitors.

Poisons completely reduce the activity of the catalyst.

Inhibitors partially suppress the activity and can change the selectivity of the catalysts.

Solid catalysts on the outer and inner (inside the pores) surface have active centers - polyhedra. These active sites on the surface are distributed by energy (activated adsorption and chemisorption), and their share of the total catalyst surface ranges from 1 to 10 percent. On this basis, a very small amount of poison is required to completely poison the catalyst, i.e., to reduce their catalytic activity to zero.

Loss of activity occurs due to partial or complete shutdown of the active surface of the catalyst. The mechanism of poisoning is specific for a given poison and catalyst and is diverse. The action of the poison can be selective, which makes it possible to increase the selectivity of the catalyst.

Consequently, the poison molecules, when adsorbed on the surface of the catalyst, do not cover the entire surface of the catalyst, but only its active part, active centers, thereby poisoning the effect of the entire catalyst for the chemical process.

The resistance of a catalyst to the action of contact poisons is the most important criterion for its applicability in production. Often, of several proposed catalysts, less active, but more resistant to poisoning, are adopted for operation.

When poisoning contact masses, a distinction is made between true poisoning (irreversible, reversible, cumulative and favorable) and deactivation as a result of blocking and sintering.

6.2.1 True poisoning.

This type of poisoning occurs during the chemical interaction of the poison with the catalyst with the formation of a catalytically inactive compound or as a result of activated adsorption of the poison on the inactive sites of the catalyst.

In chemical poisoning, the activation energy increases.

In the case of adsorption poisoning of an inhomogeneous surface, the activation energy can increase gradually. This may change the order of the reaction. Adsorption poisoning of a homogeneous catalyst is not accompanied by a change in the true activation energy, and the observed activity depends linearly on the poison concentration on the catalyst surface.

Poisoning can be reversible, irreversible, cumulative, favorable.

Reversible catalyst poisoning

With reversible poisoning, the activity of the catalyst decreases to a certain level corresponding to the concentration of the toxic impurity, and then, with a further increase in the time of poisoning, remains unchanged. When the supply of the poisonous substance to the reaction mixture is stopped and the reaction mixture is replaced with fresh raw materials that do not contain poison, the activity of the catalyst is quickly restored.

Reversible poisoning of the catalyst by poisons occurs with moderate binding of poison molecules to active sites or catalyst molecules.

Reversible poisoning Pt -catalyst is carried out by CO molecules in the reaction of hydrogenation of benzene with hydrogen. When CO is supplied to the flow of the reaction mixture consisting of C 6 H 6 and H 2 , the activity of the catalyst gradually decreases. After the supply of CO to the reaction mixture is stopped, the activity of the catalyst returns to its original value after some time. This is due to the displacement of CO molecules from the active centers by hydrogen and C 6 H 6 and washing them out of the reaction volume.

Irreversible poisoning of catalysts occurs during the chemical interaction of poison molecules or ions with the active centers of the catalyst to form stable inactive compounds. Poisons of metallic platinum used in the hydrogenation of cyclohexene or decomposition of H 2 0 2 , are mercury, lead, bismuth, tin. Toxic to platinum Cu + , Ag + , Zn 2+ , Cd 2+ , Hg 2+ , In 3+ , Ti 3+ , Co 2+ , Fe 2+ .

With irreversible poisoning, the activity of the catalyst sharply decreases. Substances that irreversibly poison the catalyst must not be used in its manufacture. Especially one has to be wary of such typical poisons (for a number of processes) as compounds of sulfur, phosphorus, arsenic, etc. Hydrogenation catalysts of the type Ni, Pt and Pd poisoned by sulfur compounds.

The poisoning effect of the poison depends on the operating temperature of the catalyst and the pressure in the reactor, on the nature and method of preparation of the catalyst. Catalyst poisoning does not occur at T = 973 K and above, since at this temperature catalysts often lose their full catalytic activity due to structural changes, and volatile poisons can be completely destroyed. The degree of catalyst poisoning depends on the composition and structure of polyhedra in the lattice of solid catalysts. This is reflected in the sensitivity of the catalysts to the poison.

Cumulative (accumulating) poisoning is expressed in the progressive deactivation of catalysts under the action of small amounts of poisons contained in the reagents.

Cumulative poisoning of catalysts occurs with the slow accumulation of a poisonous substance on the catalyst (on the outer and inner surfaces) in the course of the corresponding reactions. Poison molecules can accumulate on the catalyst by running side reactions along with the target reaction or by gradually removing the poison molecules from the reaction mixture.

An example of cumulative poisoning of catalysts is the accumulation of coke deposits in the processes of conversion of hydrocarbon fractions during the cracking of oil fractions on alumina zeolite silicate catalysts; hydrodesulfurization of oil fractions on aluminum-cobalt-molybdenum catalysts; when reforming gasoline on platinum-rhenium catalysts on aluminum oxide.

Coke gradually accumulates on the catalysts, reducing their activity, and the more coke is deposited on the catalyst, the lower its activity. However, the activity of the catalyst can be restored, if not by 100%, then by 85–90% after coke is burned off the surface of the coked catalyst in an air stream at temperatures above 773 K.

During cracking of heavy oil fractions on zeolite-aluminum-silicate catalysts, nickel and vanadium compounds can be deposited on their surface, the oxidation of which in the air stream during the burning of coke produces oxides NiO, V 2 O 5 and FeO . These oxides are deposited on the surface of the catalyst and irreversibly reduce the activity of the catalyst in cracking. Irreversibly reduce the activity of the catalyst water vapor, which at elevated temperatures have a negative effect on the texture of the catalyst.

The mechanism of poisoning is associated with the chemical composition of the catalysts and, accordingly, the type of catalysis; it will be different for electron (homolytic) catalysis on semiconductors and metals and ionic (heterolytic) catalysis. The mechanism of poisoning on metal and semiconductor contacts is the most complicated. Semiconductor type catalysts are more resistant to poisons than metal ones. The process of poisoning semiconductor contacts has been studied much less than metal ones.

Favorable poisoning of catalysts occurs when poisons introduced into the catalyst partially etch away the individual active sites of the catalysts. This ensures that the poison molecules inhibit the formation of the final reaction product or reduce the formation of intermediate reaction products. An example of favorable poisoning is the change in the selectivity of the palladium catalyst in the hydrogenation of benzoyl chloride without and with the addition of toxic substances.

The process of reduction of benzoyl chloride in boiling toluene proceeds according to the scheme:

C 6 H 6 COS1 + H 2 --> C 6 H 5 CHO + H 2 --> C 6 H 5 CH 2 OH + H 2 --> C 6 H 5 CH 3.

On a pure palladium catalyst, toluene was the final product. When quinoline was added to the mixture in an amount of 0.1 to 50 mg/kg, the process on the catalyst stopped at the stage of benzaldehyde production, which was obtained in an amount of 23 to 78–88% wt.

6.2.2 Deactivation due to blocking and sintering.

Catalyst activity can be reduced not only by true poisoning, but also due to changes in structural characteristics, as well as mechanical screening of the catalyst surface by dust or solids formed during catalysis (blocking).

For finely porous catalysts operating at relatively low temperatures, blocking of the contact surface can occur as a result of volumetric filling of micro- and transitional pores during adsorption, capillary condensation, or precipitation of microsolid particles from the reacting mixture (for example, carbon and resins during the catalysis of reactions of organic substances).

Catalyst carbonization is observed in many processes: cracking, reforming, dehydrogenation, etc. Coke formed on the surface of catalysts always contains a certain amount of hydrogen and, in terms of chemical structure, is highly condensed aromatic hydrocarbons. The formation of coke is considered to be a side stage of the main catalytic process. According to existing data, coke is deposited on catalysts up to a certain limit.

The actual coke content depends on the temperature, the nature of the feedstock, the porous structure and the chemical composition of the catalysts.

When blocking, as a rule, neither the activation energy of the catalyst nor its selectivity (excluding processes in the diffusion region) change, since the action of the blocking substance is reduced to turning off individual sections of the active surface.

The activity of contact masses can also decrease when the porous structure changes under the influence of high temperatures (sintering).

Sintering is the aggregation of small particles into larger ones, which leads to a decrease in the active surface of the catalyst and, accordingly, to a decrease in its activity. The driving force of sintering is the difference between the thermodynamic potentials of small and large particles. Sintering, apparently, is realized by two mechanisms: due to the diffusion of particles and due to the transfer of atoms.

6.3 Regeneration of contact materials.

After a certain time, which can range from a few seconds to several years, the activity of the catalyst decreases to a level that makes continued operation economically unfeasible. It may also decrease its selectivity. In such cases, the loaded catalyst must be regenerated or replaced. Catalyst regeneration is carried out in the same catalytic reactor by running a mixture of a different composition, or the catalyst is unloaded and regeneration is carried out elsewhere.

With reversible poisoning, removal of the poison from the feedstock is sufficient for regeneration. In some cases, additional processing of the catalyst is required. For example, if Ni -catalyst is deactivated by oxygen additions, it is reduced with hydrogen, which leads to the restoration of its activity close to the original one.

Nickel, irreversibly poisoned with sulfur, when treated with water vapor or oxygen converts sulfur into an even more difficult to remove sulfate. In practice, these catalysts do not regenerate. The applied catalysts also do not regenerate after sintering.

Catalysts containing Pt, Pd, Rh and others are removed from the reactor to recover the valuable precious metal.

To regenerate the coked cracking catalyst, the flow of the reaction mixture through the catalyst is stopped and the catalyst is heated in an oxidizing atmosphere so that the coke is oxidized or "burned off".

The main task in the regeneration of a coked catalyst is to reduce the temperature rise caused by the exothermic reaction of coke oxidation to CO and CO 2 . An undesirable rise in temperature during regeneration can lead to sintering of the catalyst. One of the possible solutions is the use of low oxygen concentrations in the initial stages of regeneration.

A study of the kinetics of combustion of deposited coke showed that CO and CO are the primary products of carbon combustion. CO/CO ratio 2 almost does not depend on what form the carbon is in; it is 0.3-0.9 and increases with the process temperature. The overall rate of carbon combustion in a coked catalyst is determined either by the kinetics of carbon oxidation in the pores of the granule, or by the diffusion of oxygen in the pores of the catalyst, and also by both the kinetics of the reaction and diffusion through the pores. With an increase in temperature, there is a transition from the kinetic region of coke combustion through the transition region to the diffusion region. For cracking catalysts with a particle size of ~4 mm, the kinetic region is observed up to 475°C, the diffusion region begins at temperatures above 625°C. At low temperatures, oxygen has access to all places of coke release throughout the grain volume. At high temperatures, the process is limited by the diffusion of oxygen through the pores.

It should be noted that the regeneration not only burns out the coke, but also changes the catalyst itself under the action of the oxidizing environment. During regeneration, a catalyst can change its structure, chemical composition, porosity, and specific surface area.

If deactivation is due to precipitation of metal from impurities in the feedstock, a simple oxidation procedure is not suitable. In this case, the catalyst is unloaded from the reactor and subjected to complete processing.

Depending on the nature of the substances poisoning the catalyst or on the cause of the loss of activity, methods are being developed for the regeneration of catalysts under industrial conditions.

The regeneration of contact masses is as specific as their poisoning.

In each case, the reason for the decrease in activity and change in the selectivity of the catalyst is identified, and methods for its regeneration are developed.

Of the possible ways to restore the activity of contact masses, the most significant are the following:

1. Volatile poison can be removed from the surface of the catalyst by a stream of pure gas, liquid, or by raising the temperature.

2. during chemical interaction, the poison can turn into a non-toxic, weakly adsorbed form.

3. washing of catalysts with liquid solvents.

4. treatment with a mixture of gases that are reducing or oxidizing agents.

Example. Catalysts that have lost activity due to surface blockage during coke formation are regenerated by burning coke with air oxygen at 550-700 about S.

The catalyst can be freed from oxides of vanadium and nickel by dissolving them in aqueous solutions of inorganic acids.

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At the end of the reaction cycle, the catalyst loses its activity due to the deposition of coke on it. The regeneration process is carried out in stages. First, the acceptance of raw materials to the plant is stopped. The hydrotreating unit and the stabilization unit are switched off. The circulation of hydrogen-containing gas in the reformer continues to flush the system from hydrocarbons. Further, the fuel supply to the reformer nozzles is gradually reduced until it is completely turned off. The system is gradually cooled to a temperature of 200°C, and the circulation of the hydrogen-containing gas is stopped. Hydrogen-containing gas is discharged through pressure reducing valves into the fuel network. From the reactors, the rest of the hydrocarbon vapors are sucked off by a vacuum pump. The system is then purged with an inert gas into the atmosphere. After purging, the system is filled with inert gas to a pressure of 1 MPa, the circulating compressor is turned on, and the reactor block is gradually heated with constant inert gas circulation. At 250°C, air is added to the inert gas in such an amount that the volume content of oxygen in the inert gas does not exceed 0.5% at the beginning of regeneration and 2% at the end. Coke burning is carried out in two stages: the first stage at a temperature of 250-300°C, the second - at 380-400°C. After the end of coke burning, the catalyst is calcined at 500°C. Then the system is cooled, the circulation of the inert gas is stopped and it is discharged into the atmosphere. After that, the system is again purged with hydrogen-containing gas.

Regeneration of R-56 reforming catalyst

The regeneration of the R-56 catalyst includes the following steps:

a) stop
b) coke burning;
c) oxidation;
d) cooling (air);
e) purge with nitrogen;
f) recovery in the HSG current;
g) start.

If necessary, an additional sulfate removal step is included, which is carried out immediately before the oxidation step.

Reduce the temperature of the inlet to the R-3,4,104 reactor to 480°C.

Reduce the temperature of the inlet to R-3,4,104 to 455 °C at a rate of up to 30 °C/hour while reducing the load to 50% of the nominal.

At a temperature of 455 ° C and a load of 50%, remove the raw material, stop the supply of dichloroethane and water.

Carry out a 2-hour circulation.

Lower the temperature to 400°C, turn off the stove.

After that, the circulation is continued for 15-20 minutes, then the CK-1 is stopped and the WSG pressure is released from the system.

After purging with nitrogen, the regeneration circuit is assembled.

Raise the nitrogen pressure to 3-5 kgf/cm2, after which nitrogen circulation is established.

The temperature in the reactors is raised to 400°C, holding is carried out until the inlet and outlet temperatures are equal or stabilize.

After that, the air supply is started in such a way as to create an oxygen concentration in nitrogen from 0.5 to 0.8% vol.

The oxygen concentration is maintained such that the outlet temperature does not exceed 455 °C.

After the beginning of the air supply, proceed to the gradual supply of dichloroethane to R-3 at a rate to ensure the molar ratio "water:chloride"=20:1. Dichloroethane feed schedules to ensure proper ratio are provided by UOP Technical Consultant.

The combustion of coke is considered complete when the temperature difference remains unchanged for 4 hours at the same oxygen concentration at the inlet and outlet of the reactor. If it is planned to depressurize the system (for repair work), then it should be limited to the stage of coke burning, and the remaining stages should be completed after repair work.

The oxidation stage lasts 11 hours. At the end of the 4-hour exposure, the outlet temperature rises to 510°C within 4 hours.

The oxygen concentration is maintained at 0.6-0.8%.

When the temperature reaches 510°C and there is no combustion of coke, the oxygen concentration is gradually brought to 5%. If there is evidence of burning residual coke, then the oxygen concentration should be maintained such that in no case does the outlet temperature exceed 520°C, as burning coke at a temperature of 510°C causes much more damage to the catalyst than conventional coke burning.

In the absence of coke combustion, keep the oxidation parameters - temperature 510 ° C, oxygen concentration 5% - for 11 hours. The time is counted from the moment of stabilization of the temperatures at the outlet of their reactors. Adsorbers K-108, K-109 are in operation during catalyst regeneration.

Supply of dichloroethane should be carried out with the calculation of the ratio "water: chloride" = 20: 1.

After 8 hours of oxidation, set the supply of dichloroethane with the calculation of the ratio "water: chloride" = 40:1.

After 11 hours, stop the supply of dichloroethane, drain the liquid from all low points.

Upon completion of the oxidation, the furnace is switched off and circulation is continued until the temperature at the outlet of the reactor drops to 200°C.

Upon reaching 200 C, stop PC-3, relieve pressure.

Nitrogen is purged to an oxygen content of less than 0.5% vol.

At the end of the purge, disassemble the regeneration circuit.

Take WASH on block, drain fluid from all low points.

Build up the WSG pressure of 5-6 kgf/cm, start up PC-3, raise the reactor inlet temperature to 430°C at a rate of 40-55°C/hour.

At the time of rising temperatures, exhaustively drain the system.

The recovery lasts 1 hour or more, controlled by the concentration of H^S at the outlet of the reactors (2ppm or less).

Upon completion of the reduction, the temperature at the inlet to the reactor is lowered to 370°C and the installation is started up, as described above.

Regeneration gas alkalizing unit

A fresh 42% alkali solution from a tank truck is fed into the E-202 tank, from where it is pumped by the H-202 pump into the E-201 tank. The pump discharge pressure is controlled by pos.RIAN-428. After that, chemically purified water is supplied to the E-201 tank, the solution is mixed and a 2% NaOH solution is supplied according to the scheme:

Before the circulation of alkali begins, it is necessary to fill the protected equipment with alkali, for which alkali is fed into the gas-air mixture flow before X-106,106a, then into X-6, X-6a and C-7.

The temperature at the outlet of the refrigerator X-106, X-106a, X-6, 6a is controlled.

After the liquid level appears in the C-7 separator, the E-102 tank is turned off and the H-201 pump switches to work from the C-7 separator, after which the alkali circulation is established according to the above scheme.

Periodically, samples are taken from the C-7 separator to control the concentration of alkali in the circulating solution. When the alkali concentration drops to 1% of the mass, a fresh 2% alkali solution is fed from the E-201 tank by the H-202 pump to the N-201 pump discharge line. Simultaneously with the feed, part of the spent alkali is discharged from the separator into the E-8 tank .

Regeneration of hydrotreating catalyst AKM

After the installation is stopped, the intershop communications are turned off and the apparatus is freed from the oil product, the compressor VK-1 (VK-2) on nitrogen is put into operation.

Compressor discharge nitrogen is supplied to C-5 separator and then to T-1/1-3 heat exchangers.

After heating in the furnaces, nitrogen enters the R-1 reactor, and then into the R-2. From the R-2 reactor, nitrogen enters the tube space of the T-3 heater, then passes through the tube space of the T-1 / 1-3 heat exchangers and then to the separator with a vent to the atmosphere.

During purge, nitrogen samples are taken from the C-4 to determine the oxygen content. When the oxygen content is less than 0.5% vol. the outlet is closed, the pressure in the system increases, the circulation compressor PK-1 is switched on and nitrogen circulation is established to flush the system from combustible gases. During the "flushing" fresh nitrogen is constantly supplied to C-5, and a constant purge from C-4 is carried out.

It is possible to supply nitrogen by the VK-1 compressor of the 24-6/2 unit, which allows simultaneous regeneration of hydrotreating and reforming catalysts.

During flushing, nitrogen samples are taken to determine the concentration of combustible gases. With a decrease in the content of combustible gases to 0.5% vol. the temperature rise begins in the reactors R-1, R-2. At a temperature of 150°C, the air supply begins while controlling the temperature difference in the catalyst bed. When the combustion is weakened or stopped, the temperature in the R-1, R-2 reactors is raised to 250-300 ° C and the air supply is increased.

At the last stage of coke burning, the temperature in the reactors is raised to 400°C and the VK-1 compressor is completely transferred to air.

The end of combustion is determined by the oxygen concentration in the regeneration gases at the inlet (C-5) and outlet (C-4), which should be the same.

At the end of regeneration, the temperature in the reactors is reduced with the simultaneous replacement of regeneration gases with nitrogen.

Having extinguished the furnaces P-1, P-104, the temperature in the reactors R-1, R-2 continues to decrease to 100°C (or lower), controlling the concentration of oxygen in nitrogen. Purge is considered complete if the oxygen content in nitrogen is less than 0.5% vol.

Carry out an internal inspection of the reactors, remove foreign objects, clean the gutters and the grid of the central pipe with a vacuum cleaner. Especially carefully check the presence of plugs on the pockets of the troughs, punch all the slots around the central pipe, along the perimeter of the reactor with an asbestos cord.

Catalyst loading using the UOPi loading machine can be carried out in dry, clear weather. During rain, snowfall, loading is prohibited. To ensure the operation of the loading machine, dried air with a pressure of 3-4 kgf/cm2 must be supplied to the reactors. The use of nitrogen is not recommended as when mounting the upper plate, it is necessary to go down into the reactor.

Loading should be carried out at the rate recommended by the USPE specialist, who is usually present at the time of loading, or by reference to the loading charts provided by USPE.

At the end of the loading, it is necessary to assemble the upper plate, remove the plugs from the grooves and seal all the gaps between the plate segments and the support ring with an asbestos cord. After that, the reactor should be closed and the installation start-up procedure should be started.

Recovery of reforming catalysts

In the process of recovery from the initial oxidation state of Pt +4 and Re +7, these elements are reduced to metals.

Recovery of APC leads to a decrease in its activity. This is due to a decrease in the dispersion of platinum. If the calcined APC is not stored in a sealed container, then it absorbs water from the air (up to 10% wt.). Recovery of the rehydrated catalyst also leads to a decrease in the dispersion of platinum.

The reason for these phenomena is that H 2 O weakens the interaction of PtO 2 and Al 2 O 3 . Platinum becomes more mobile and sintering occurs.

It is recommended to reduce the APC with dry hydrogen at a minimum pressure (0.1 MPa) and a temperature of up to 500 0 C. Under these conditions of reduction, the size of the crystallites is about 1.0 nm. With increasing pressure and temperature, the size of the crystallites increases.

2.3.4. Sulfurization of reforming catalysts

APC (AP-56, AP-64) was sulfurized by treating wet granules with hydrogen sulfide after impregnation with H 2 PtCl 6 . Aluminum-platinum-rhenium catalysts are also subjected to sulphurization, but after recovery.

This operation is necessary to suppress the excess activity of catalysts in hydrogenolysis, which leads to a drop in the hydrogen concentration in CVSG and contributes to premature catalyst coking.

As a result of strictly dosed sulfurization, the rate of dehydrogenation of six-membered naphthenes on the Pt-Re/Al 2 O 3 catalyst decreases by 2.5 times, and the rate of paraffin hydrogenolysis, by 400 times.

For Pt-Sn/Al 2 O 3 and Pt-Ge/Al 2 O 3 catalysts, sulfiding is not required, because Sn and Ge inhibit hydrogenolysis.

2.3.5. Sintering and redispersion of platinum.

The operating life of modern reforming catalysts is 5–10 years. Service life before the first regeneration up to 5 years. After regeneration, the duration of the reaction period, as a rule, is reduced.

The reasons for this are as follows.

In the process of oxidative regeneration, coke deposits are oxidized:

C m H n S x + O 2 ® CO 2 + H 2 O + SO 2

2 SO 2 + O 2 ® 2 SO 3

SO 3 alone or through the formation of H 2 SO 4 reacts with alumina:

3SO 3 + Al 2 O 3 ® Al 2 (SO 4) 3

Al 2 (SO 4) 3 decomposes only at temperatures above 750 0 C. Therefore, sulfate sulfur accumulates in the reformin catalyst during oxidative regeneration. The spent reforming catalyst may contain 0.1 - 0.8% wt. sulfate sulfur (Kozlov).

A catalyst with such an amount of sulfate sulfur completely loses its activity. Sulphates inhibit the oxidation of coke. Basically, when coke is oxidized, it is not CO 2 that is formed, but CO. Carbon monoxide on a chlorinated catalyst forms volatile platinum carbonyl chlorides. This leads to coarsening of platinum crystallites.

After burning out the coke, platinum is redispersed by oxychlorination. The catalyst is treated with a gas mixture containing C1 2 (or its compounds: HC1, paraffin chlorine derivatives), oxygen, and water vapor. The processing temperature is below the temperature of oxidative regeneration - 450 0 C. The reaction gas mixture contains C1 2 and HC1. At a temperature of 450 0 C, a volatile platinum compound is formed - PtCl 2. Under these temperature conditions, the formation of OH groups on the surface of aluminum oxide is possible, to which platinum can be attached. Thus, it is dispersed. Then the catalyst is calcined in air at 580 0 C, in which case the oxidation of platinum Pt +2 ® Pt +4 occurs, and the platinum is stabilized by the support.

2.4. Industrial reforming catalysts

Axens has in the past offered both the now traditional pratinorhenium catalysts (RG 482) and the trimetallic RG 582. It is reported that the introduction of a third metal into the catalyst structure increases the yield at the beginning of the C 5+ cycle by approximately 1%, while the cycle time is comparable to performance of a standard Pt/Re catalyst.

Compared to this catalyst, the new RG 682 offered by the company increases the C 5+ yield by 0.6 - 1%. As a moderately unbalanced catalyst (i.e. containing more Pt than Re), RG 682 is characterized by greater resistance to coke attack, and as a result, its cycle time is more than 35% higher than that of the "classic" balanced bimetallic Pt / Re catalyst.

Modern trimetallic reforming catalysts differ from Pt/Re in reduced hydrogenolysis, which increases the yield of stable hydrogenate and makes it possible to obtain purer HSG.

Nanotechnology techniques were used in the development of this catalyst.

Due to the lower sulfur content of the 500 Series Catalysts, they are more tolerant of sulfur slippage. The use of 600 series catalysts is preferred when low sulfur feedstocks are to be processed.

Table 2.1

Content, % wt.	RG 582 A	RG582	RG 682 A	RG 682
Pt	0,27	0,3	0,27	0,3
Re	0,27	0,3	0,4	0,4

Lecture 13

3.Catalysts for catalytic cracking

Catalytic cracking began to develop in the 1940s, when the first installations with a fixed catalyst bed were created, which were used as granules of acid-activated bentonite. In the 1950s, catalysts based on natural clays were replaced by amorphous synthetic aluminosilicates. These catalysts were much more active than those based on clays and had greater stability and mechanical strength. The creation of synthetic aluminosilicate catalysts made it possible to move from the technology of cracking with a stationary bed to the technology with a moving catalyst bed, and later to the technology with a fluidized bed of a microspherical catalyst. The development of fluidized bed technology required an increase in the mechanical strength and activity of catalysts and a decrease in coke formation. This problem was solved in the mid-1960s by the creation of amorphous aluminosilicate catalysts with a high content of AI 2 O 3 . At the same time, zeolite cracking catalysts were developed, which are several orders of magnitude more active than amorphous aluminosilicate catalysts.

The subsequent improvement of zeolite catalysts (increase in activity, stability and selectivity) led to the development of a new technology for their use - the process of cracking with a moving flow of a pulverized catalyst (cracking in a riser reactor), which was implemented in industry in the early 80s and is currently is a leader in oil refining.

Catalysts of the first generation (1970-1978) were intended for the processing of relatively light and noble raw materials - kerosene-gas oil fractions and light vacuum gas oils, they were systems with a highly developed specific surface area, a large pore volume, and a low zeolite content (8-12%) . The content of Al 2 O 3 . ranged from 10-13% (low alumina) to 25-30% (high alumina). Catalysts of the second generation, created in the period 1976–1982, were distinguished by an increased content (up to 15–20%) of zeolite in a deeply substituted cationic (REE) form, a catalytically inert matrix with a low specific surface area. The content of A1 2 O 3 in these catalysts was 30-35%. Due to the high content and microporous matrix, these catalysts had a sufficiently high thermal and thermocouple stability, increased coke selectivity, and relatively low metal resistance, up to 2000–3000 ppm.

The development and implementation of third-generation cracking catalysts was facilitated by changes in the technology of cracking (once-through lift-reactor with a forced fluidized bed) and regeneration (high-temperature regeneration with post-burning CO), as well as a further increase in the requirements for thermal and thermocouple stability, increased selectivity for coke and gas, greater metal resistance and obtaining the maximum yield of gasoline.

Vacuum gas oil cracking catalysts are high-alumina (A1 2 O 3 content is 40-45%), contain a wide-pore aluminosilicate matrix with moderate catalytic activity and 20-30% RZU zeolite with a molar ratio of SiO 2 /Al 2 O 3 = 4.6-5 .0 and low content of residual Na 2 O (0.2-0.5%). Third generation catalysts are characterized by high bulk density (800-1000 kg/m 3 ) and abrasion resistance, uniform granulometric composition.

1) B During operation, the catalyst gradually loses its activity as a result of coking and deposition of raw materials metals on its surface.

2) D To restore the original activity, the catalyst is subjected to regeneration by oxidative burning of coke.

3) B Depending on the composition of the catalyst, a gas-air or steam-air regeneration method is used. Zeolite-containing hydrodesulfurization and hydrocracking catalysts cannot be subjected to steam-air regeneration.

4) G Azo-air regeneration is usually carried out with a mixture of inert gas and air at temperatures up to 530 °C. In this case, the regenerated catalyst accelerates the coke combustion reactions.

5) P Air-air regeneration is carried out with a mixture heated in a furnace to the temperature at which coke starts to burn out. The mixture enters the reactor, where layer-by-layer burning of coke takes place, after which the gases are discharged into the chimney.

4 Description of installation year

Hydrotreating is the deepest form of hydrogenation processes. Both straight-run distillates (gasoline, kerosene, diesel fuel, vacuum gas oil) and secondary distillates (light fraction of pyrolysis tar, gasolines, light gas oils of coking, catalytic cracking, thermal cracking, visbreaking) are subjected to hydrotreating. Hydrotreating is used to remove sulfur, nitrogen, oxygen-containing compounds from raw materials, as well as to hydrogenate unsaturated hydrocarbons.

The feedstock (diesel fuel) is preheated in heat exchangers (not shown in the diagram), mixed with circulating HSG, and fed into furnace 3, where it is heated to a temperature of 380-400 0 С (depending on the type of feedstock). After the furnace, the mixture enters the reactor 4. In some installations, 2 or 3 stage cleaning of the raw material is often provided. To increase the temperature of the mixture or remove the exothermic effect of the reaction between the reactors, cold HSG is usually introduced. After the last reactor, the hydrogenated product enters the high-pressure gas separator 6, where the flashing process usually takes place at a pressure equal to or slightly lower than the pressure in the reactor. The temperature in the gas separator is 80-85 0 С. By selecting the temperature in the separator, the concentration of hydrogen (Н 2) in the circulating HSG is controlled. The gas phase enters the absorber 8, where hydrogen sulfide (H 2 S) is captured by aqueous solutions of mono-, diethanolamine. After cleaning, a part of the circulating HSG is removed from the plant in the form of an exhaust, and the main part is replenished with fresh HSG.

After the high-pressure gas separator 6, the hydrogenated product enters the low-pressure gas separator 7, where a gas phase appears due to a decrease in pressure. The gas phase enters the absorber 9, where it is purified from hydrogen sulfide, and is discharged from above into the dry gas line.

The hydrogenate from the low-pressure gas separator 7 enters the fractionating absorber 12, where dissolved gases are removed from the diesel fuel, which are fed to the absorber 10 for hydrogen sulfide purification, and the gasoline fraction. The gasoline fraction is used as irrigation 12, and its balance amount is pumped out of the plant. Hydrotreated diesel fuel is discharged from the bottom 12, part of which is used as a hot jet of the bottom of the column 12, heated in the furnace 13. Gasoline is used as the absorbent of the column 12.

In desorber 11, regeneration of absorbent streams (monoethanolamine) saturated with hydrogen sulfide occurs in parallel. Hydrogen sulfide is discharged from above 11, and from below the regenerated absorbent is fed into absorbers 8, 9, 10.

1 - distributor; 2 - fitting for thermocouple; 3 - top bottom; 4 - casing; 5 - body; 6 - plate; 7 - lining; 8 - gutter; 9 - catalyst; 10 - central pipe; 11 - support belt; 12 - support; 13 - bottom bottom; 14 - porcelain balls; I - input of raw materials; II - output of the product; II - catalyst output

Figure - Hydrotreating reactor

The main apparatus of hydrogenation plants is a reactor with a fixed catalyst bed.

The two-section diesel fuel hydrotreating reactor is a vertical cylindrical apparatus with elliptical bottoms.

The upper layer of the catalyst is poured onto the grate, and the lower layer onto porcelain balls, which fill the spherical part of the lower bottom. To remove excess reaction heat, a collector for supplying cold HSG is installed under the grate.

The raw material supplied through the fitting in the upper bottom is evenly distributed over the entire section and first, to retain mechanical impurities, they pass

through filtering devices consisting of mesh baskets immersed in the top layer of the catalyst.

The gaps between the baskets are filled with porcelain balls. The feed gas mixture is passed through the catalyst bed in both sections and through the nozzle

the lower section is removed from the reactor.

Material balance of the GO process of various fuels