FCC PROCESS FOR MAXIMIZING DIESEL

A process is described for maximizing the FCC middle distillates comprising the use of two different converters, operating in a coordinated manner that seeks to maximize the production of LCO for diesel, generating a specified gasoline and reducing fuel oil production. Converter “A” operates with a low contact time in the riser, of 0.2 to 1.5 sec. (preferably from 0.5 to 1.0 sec.) making a higher reaction temperature possible even at low severity, from 510° C. to 560° C. (preferably from 530° C. to 550° C.) and with a catalyst suitable to the maximization of LCO. Converter “B” possesses a high activity catalytic system, suited to cracking naphtha and DO generated in the first converter. Preferably, converter “B” has two separate risers, allowing the reaction temperatures of each to be adjusted independently according to the range most recommended for maximizing the cracking of each of the streams: 530° C. to 560° C. for the DO riser and 540° C. to 600° C. for the naphtha riser. The high-quality LCO stream generated by cracking at low severity in converter “A” is not contaminated by the poorer quality LCO generated by re-cracking the DO in converter “B,” since each converter has its own fractionating tower. The use of low contact time as a route for reducing severity in converter “A” geared towards the production of better quality LCO allows it to operate with a higher reaction temperature for the same LCO conversion and quality level, entailing greater operating reliability for the unit, and providing benefits for the heat balance of the converter. In existing units, the improvement in the heat balance provides leeway to the air blower via increased batch temperature, and makes room for processing more residual batches.

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Description
FIELD OF THE INVENTION

The present invention relates to the field of fluid catalytic cracking processes (FCC) and concerns the maximization of middle distillates. More specifically, this invention describes an FCC process using heavy hydrocarbons or mixture of hydrocarbons as feedstock, which proposes the use of two different converters operating in a coordinated manner, one of the converters being dedicated to the production of top quality LCO (Light Cycle Oil) and another converter dedicated to “re-cracking” unwanted (decanted oil) or non-specified (naphtha) streams generated in the first converter. Moreover, the invention describes the benefit of using low contact time as a method for reducing the severity of the reaction in the converter, so as to produce quality LCO. The process seeks both to maximize the production of LCO and to improve its quality, as it makes it possible to produce gasoline with the required quality and to reduce fuel oil production, eliminating the typical problems of a maximum LCO operation.

THE BASIS OF THE INVENTION

As a result of the higher efficiency of diesel engines versus gasoline engines, coupled with the increased consumption of gasoline's competitors, such as alcohol and natural gas, the worldwide demand for diesel is growing at a higher rate than that for gasoline. Consequently, refiners throughout the world are interested in maximizing their diesel production units.

One of the streams produced in FCC units is the LCO (Light Cycle Oil), which has a boiling range similar to that of diesel. Its addition to the diesel pool is primarily limited by its high aromaticity and density. Improvements in the quality of this stream in general enable its percentage in the diesel pool to be increased, thus enabling refineries to produce more diesel.

Maximizing the production of middle distillates with FCC is not in itself a novelty. The solution most adopted for maximizing LCO production involves minimizing the severity of the reactions. The most common practice relates to operating the converter at a low reaction temperature (TRX) of about 450° C. to 500° C., associated with the use of a low activity catalyst. LCO production increases and its quality improves in these conditions. The problem is that unwanted factors are associated with this benefit, such as increased production of Decanted Oil (DO) and lower quality cracked naphtha, thus preventing its introduction into the gasoline pool. Implementing the FCC operation at very low reaction temperatures likewise reduces the operating reliability of the unit, in addition to its negative impacts on the converter heat balance.

The proposed invention relates to a fluid catalytic cracking (FCC) process comprising two converters operating in a coordinated manner. Converter here and throughout the text refers to an assemblage of equipment enabling: cracking reactions of the feedstock, including the use of finely particulate catalyst enabling a catalytic reaction mechanism; the regeneration of the spent catalyst; and the transport of the fluidized catalyst with rheological characteristics of a powder, between reaction zone and regeneration zone. In addition, the converter itself includes a single catalyst formulation in its inventory.

The proposed process seeks both to maximize the production of LCO and to improve its quality and, in addition, to produce gasoline with the required quality and reduce fuel oil production, thus eliminating the typical problems of a maximum LCO operation.

CORRELATED TECHNIQUES

Catalytic cracking in a fluidized bed plays vital role in oil refining, particularly when heavy, therefore, difficult to distill hydrocarbons must be processed. The main characteristics of the process are: the possibility to adjust the production according to actual market needs and reuse of low commercial value fractions coming from other refinery processes.

The FCC unit has intermediate and heavy oil fractions, generating higher added value, lighter products (lighter gasoline and intermediates) through chemical reactions breaking down the molecules via zeolitic catalysts. This type of cracking takes place at controlled temperatures that are lower than thermal cracking.

The FCC process has been used industrially since the 1940s and is therefore a technique of the process that is amply described in literature. An introductory reference to the FCC process can be found in the publication: Fluid Catalytic Cracking Technology and Operation; Wilson, J. W; Ed. Pennwell Books; 1997.

A summary of the basic aspects of the FCC technology can be found in Patent Application PI 0504321-2, used as a reference for the below description:

Cracking reactions occur due to the contact of hydrocarbons in a tubular or riser reaction zone, with a catalyst comprising fine particulate material. The feedstock most commonly subjected to the FCC process are, in general, those streams of oil refineries coming from side-cuts of vacuum towers, known as heavy vacuum gas oil (HVGO), or heavier cuts like the ones coming from the bottom of atmospheric towers, called atmospheric residue (AR), or else mixtures of these streams.

These streams, which have a density typically in the range of 8° to 28° API must undergo a chemical process, such as the catalytic cracking process, which essentially alters their composition, converting them into higher economic value, lighter hydrocarbon streams.

During the cracking reaction, substantial proportions of coke, a by-product of the reaction, are deposited on the catalyst. Coke is a high molecular weight material, consisting of hydrocarbons typically containing 4% to 9% by weight of hydrogen in its composition.

At the end of the reactions in the riser, the recovered coke catalyst, generally known as “spent catalyst,” must be separated from the cracking products. This separation operation is typically done in cyclones contained within a vessel known as a “separator vessel.”

The Applicant's separation technology, covered in the patent documents PI 9303773-2 (which includes a rapid separation system at the end of the riser), PI 0204737-3 (which includes an improved rapid separation system by introducing hydrocarbon vapor collecting tubes into the system) and PI 0704443-7 (which includes the application of the rapid separation system for converters with two or more risers), fully incorporated herein as a reference, is the preferred technique for use in this process.

The separated spent catalyst still contains some amount of noble products from cracking entrained among the particles of the catalyst or absorbed into its surface and must, therefore, be rectified to recover this residual hydrocarbon. This operation takes place in the rectifier, equipment generally coupled to the separator vessel. The rectification fluid is, generally, steam, and this steam flows counter to the flow of spent catalyst. The internals of the rectifier permit closer contact among the catalyst particles and the rectification steam.

The extracted hydrocarbons and the rectification steam join the cracking product stream, previously separated in the cyclones, and allow the separator vessel to continue through the transfer line up to the fractionating tower, where the characteristic streams of the FCC are recovered:

FG: fuel gas—inerts of the process, H2S, H2 and C1 and C2 hydrocarbons;

LPG: liquified petroleum gas C3 and C4;

Acidic waters: the sum of all the vapor flows injected into the reaction zone and into the top system of the fractionation section, which contains high levels of contaminants such as H2S;

GLN: cracked naphtha specified as gasoline—hydrocarbons C5+ final boiling point (FBP) of 220° C., typically;

LCO: light cycle oil—hydrocarbons with an IBP of 220° C. up to an FBP of 340° C., typically;

DO: decanted oil—hydrocarbons with an IBP of 340° C.+ (residual product), typically.

The cracked naphtha stream can be fractionated into light naphtha (LN) and heavy naphtha (HN), being at the cutting temperature most suitable to the refiner. In addition, in some configurations, a circulating reflux stream is removed from the tower as an intermediate product between the LCO and the DO, heavy cycle oil (HCO) stream

The spent catalyst at the outlet of the rectifier containing non-rectifiable carbonaceous material is subsequently sent to a regenerator vessel. The coke that is deposited on the surface and in the pores of the catalyst is then burned in this vessel, which operates at high temperature. The elimination of coke through combustion allows the recovery of the catalyst activity and releases heat in a sufficient quantity to meet the thermal needs of the catalytic cracking reactions.

The fluidization of the catalyst particles by gaseous streams enables the transport of catalyst between the reaction zone and the regeneration zone and vice-versa. Apart from fulfilling its essential function of promoting the catalysis of the chemical reactions, it also acts as a means of transporting heat from the regenerator to the reaction zone.

There are many descriptions of this technique involving hydrocarbon catalytic cracking processes in a fluidized catalyst stream, transporting the catalyst between the reaction zone and the regeneration zone and burning coke in the regenerator.

Since its initial conception, the fluid catalytic cracking (FCC) has essentially been geared towards producing high-octane gasoline, and is also responsible for producing LPG. The middle distillate (LCO) produced in this conventional process represents 15% to 25% by weight of the total yield. Normally the LCO has a high concentration of aromatics, reaching over 80% of the total composition, a fact that makes its incorporation into the diesel oil pool difficult.

The aromatics in the LCO range impair its quality, reducing the cetane number (index based on the linear paraffin hydrocarbons with the chemical formula C16H34 used as the standard for assessing the ignition properties of the diesel) and increasing the characteristic density, which cannot be completely reversed by the hydrotreatment. Minimize the formation of aromatics in the FCC and, therefore, an important strategy for maximizing the middle distillates. The aromatics in the FCC products can come directly from the feedstock through dealkylation reactions of larger aromatic molecules or can be formed from olefins that cyclize and then undergo hydrogen transfer reactions. The formation of aromatics from olefins occurs in secondary reactions and the low reactivity of the aromatics causes these molecules to end up concentrating in the products as the conversion progresses and removes other more reactive types from the reaction medium. The secondary aromatic forming reactions, undesirable for obtaining a better quality middle distillate, are favored by high severity conditions including: high reaction temperature and high contact temperature between the batch and the catalyst in the riser.

Despite the fact that the FCC process was originally developed for the production of high-octane gasoline, the inherent flexibility of the process allows it to be adapted to different production objectives, such as for example maximizing middle distillates for diesel oil or even light olefins (present in FG and LPG) for the petrochemical industry. The maximization of middle distillates and the maximization of light olefins are diametrically opposed types of operation. The operation for middle distillates seeks low conversions, so as to preserve the LCO yield (boiling point range from 220° C.-340° C.), which is greater at conversion levels below those usually used in the typical FCC operation. The operation for maximizing light olefins, on the other hand, seeks very high conversions entering into the area of over-cracking of gasoline. For this reason, the task of simultaneously obtaining high olefin and better quality LCO yields in a single reaction zone is difficult.

In view of the increasing demand for high quality middle distillates, to the detriment of the market for gasoline, changes were introduced in the method of operation of FCC units, in order to increase the production of LCO in this process. Several documents in the patent literature discuss changes in the catalytic systems and to the process variables, in order to achieve reduced process severity, so as to increase the yield of middle distillates and to reduce the aromatic content in this fraction. Among these changes are:

    • reduced reaction temperature (TRX);
    • reduced contact time;
    • reduced catalyst/oil (C/O) ratio; and
    • reduced catalytic activity.

All these modifications in the operational variables lead to a reduced conversion, with the consequent increase in the production of DO and a worsening of the quality of the cracked naphtha.

The most commonly used solution for maximizing the production of LCO involves minimizing the severity of the reactions by operating the converter at a low reaction temperature, between 450° C. and 500° C., combined with the use of a low activity catalyst. LCO production increases and its quality improves in these conditions. The problem is that when this solution is applied, some undesirable side effects occurs, such as increased production of DO and lower quality cracked naphtha, what prevents the adition of that naphtha into the gasoline pool. Generally, another associated undesirable factor caused by the reduction of the reaction severity is reduced production of LPG and, consequently, of light olefins in the LPG.

A negative consequence resulting from the FCC operation for maximization of middle distillates by the reduction of the reaction temperature, is the excessive deposition of coke observed in the reaction zone and in the rectification zone. The coke deposition in the riser, cyclones of the separator vessel, transfer line and bottom of the rectifier is commonly seen in low temperature reaction conditions. This problem reduces the operating reliability of the unit and often requires its complete shutdown for removal and cleaning of the coke.

The primary cause of the excessive deposition of coke is the incomplete vaporization of the feedstock, due to the low resultant temperature in the mixing zone between the feedstock and the catalyst, as a result of the lower reaction temperatures used for maximizing the middle distillates.

Another negative consequence of the FCC operation for maximizing middle distillates, by the reduction of the reaction temperature, is the fast rising of the regenerator temperature, the operating range of which must be between 670° C. and 730° C. in the dense catalyst bed (dense phase temperature of the regenerator). The excessive temperature rise of the regenerator is detrimental, because it increases the deactivation rate of the catalyst inventory of the unit, in addition to bringing the regenerator vessel close to its metallurgical design limits. Furthermore, it has the undesirable effect of significantly increasing the potential gum of the cracked naphtha, resulting in greater thermal cracking to which the feedstock is subjected upon mixing with the hotter catalyst resulting from a hotter regeneration zone, despite the lower TRX.

The fast rising of the regenerator temperature during the reduced reaction temperature operation is the main cause of the increased entrainment of hydrocarbons to the regenerator. This additional contribution of fuel to be burned in the regenerator, associated with the reduced circulation of catalyst of the unit (due to the reduced reaction temperature in the riser), causes the fast rise of the regenerator temperature. The hydrocarbon entrainment increases, in turn, due to lower rectification efficiency, caused both by low temperatures in the rectifier, accompanying the reaction temperature, and by a higher concentration of the higher molecular weight components in the gases entering the rectifier, due to the lower feedstock conversion and the greater unvaporized portion of it, as a result of the lower temperature of the mixture at the base of the riser, accompanying the reaction temperature.

The deleterious effect of the FCC operation with a reduced reaction temperature on the regenerator is worse when processing heavy and residual feedstocks, which tend to be more difficult to vaporize and convert.

The temperature of the regenerator can be corrected by a catalyst cooler unit (catalyst cooler, or catcooler) in the regenerator vessel. The catcooler receives a stream of hot catalyst from the vessel; exchanges heat with the tube bundles through which water flows and into which saturated steam is generated, and returns a cooled catalyst stream to the regenerator, while maintaining the dense phase temperature of the regenerator in the proper operating range. Aside from increasing the complexity of the equipment, its use has the drawback of increasing the air demand to the regeneration zone, due to the increased total coke yield, resulting from its effect on the overall energy balance of the converter.

Another method for correcting the temperature of the regenerator consists of injecting a cooling “quenching” fluid stream into the riser, preferably at a point above the main feedstock injection point, as described in the patent application PI 0504854-0.

The injection of a “quench” leads to an increased energy demand for the riser, entailing increased catalyst circulation from the regenerator to the riser. This additional removal of hot catalyst from the regenerator indirectly contributes to reducing the temperature of the dense catalyst bed of the regenerator. Another benefit of the “quench” is the increased temperature at the base of the riser (for the same final reaction temperature), which contributes to the better vaporization of the feedstock (particularly important for heavy feedstock) and to its bottom cracking, reducing the production DO and also reducing coke deposition on the converter due to the presence of non-vaporized feedstock. Like the catalyst cooler, this alternative has the drawback of increasing the total coke yield of the unit, consequently requiring the use of greater combustion air-flow rate.

The above techniques do not eliminate the problem of hydrocarbon entrainment to the regenerator. In this context of a reduced reaction temperature, the Applicant's quick separation technology, the object of patent PI 9303773-2 and of the aforementioned patent applications PI 0204737-3 and PI 0704443-7 and incorporated in their entirety as reference, plays an important role in reducing the hydrocarbon batch to the rectifier, thus contributing to reducing entrainment to the regenerator.

Apart from reducing the reaction temperature, another way of reducing the reaction severity for the operation for maximizing middle distillates is to reduce the contact time between the feedstock and the catalyst in the riser. With the low contact time, secondary reactions forming aromatics are eliminated and a better quality LCO is produced. At the same time, however, the bottom cracking reactions are extended, which also entails an undesirable increased production of DO, similar to that occurring when the reaction temperature is reduced.

However, in the operation oriented towards middle distillates, the reduced reaction severity via low contact time offers advantages over that via reduced reaction temperature. This is because in opting for the use of the low contact time, it is possible to achieve the same objectives of maximizing middle distillates by operating at a reaction temperature higher than that if the reduction in severity were exclusively via reaction temperature. During the inventive process, the Applicant learned that the use of low contact temperature associated with operation with higher reaction temperature maintains the production and the improved quality of the LCO and eliminates the problems of loss of efficiency in the rectifier. In addition, the highest temperature at the base of the riser (determined by the highest TRX) allows better vaporization of the feedstock, eliminating the disadvantages of deposition of coke onto equipments of the reaction and rectification section. In the same way as the use of the “quench,” the better vaporization of the batch and at the highest temperature throughout the riser at the low contact time make possible the use of this route for cracking of residual batches. The resultant dense phase temperature is reduced both by increasing the efficiency of the rectifier and by the lower delta-coke resulting from the reduced contact time in the reaction zone, and further by a lower deposition of heavy fractions onto the catalyst, resulting in better vaporization of the feedstock, as mentioned.

For those accustomed to the FCC technology, delta-coke is the ratio of the yield by weight of coke of the feedstock to mass ratio between the catalyst circulation/feedstock flow rate (catalyst/oil, or C/O ratio). The C/O translates into unit of catalyst mass per unit of batch mass. Thus, the delta-coke indicates the quantity of catalyst that would be necessary to produce a certain quantity of coke defined by the thermal demand of the converter. The lower the delta-coke value, the greater the catalyst circulation (higher C/O) must be to generate a given coke yield in the unit, which ultimately means that: the smaller delta coke cools the regenerator (lower dense phase temperature) since it promotes the increasing in the catalyst circulation, which removes heat from the regenerator to the riser. Reducing the contact time leads to a reduced delta-coke because, under these conditions, the coke yield required to meet the thermal balance of the converter is achieved only from a greater catalyst circulation (higher C/O), which compensates the lower reaction time available in the riser.

The Applicant's patent application PI 0504321-2 presents a process aimed at eliminating the disadvantages of increased DO production and loss of the gasoline specification due to operation at reduced severity. The process is characterized by the use of a single converter with two distinct reaction sections associated with the two also distinct fractionating sections. The feedstock (stream 10) is introduced in the converter in a less severe reaction section, in a riser at a temperature of 460° C. to 520° C. (maximum LCO mode). The reaction products are separated in a fractionator from which the “LCO for Diesel” (stream 66) stream and the non-specified cracked naphtha streams (stream 64) and DO (stream 68, in high yield, due to low severity) are removed. These latter two undesirable streams are again recycled to the converter, to the other reaction section, where they are re-cracked with greater severity in a riser at a temperature from 550° C. to 620° C. The separator vessel receiving the products from this riser is independent of the previous one, which makes it possible to segregate products of this latter to another fractionator, where the specified cracked naphtha stream (stream 74), the lower yield DO stream (stream 76), and the low quality LCO (stream 75) with high aromatic content are removed. With this design, it is possible to segregate the better quality LCO from the poorer quality LCO, thus avoiding the “contamination” of these streams that have the same boiling range. Moreover, it is possible to produce naphtha specified for the gasoline pool and reduce the production of DO. The rise in the regenerator temperature is prevented by the “quench” function that the re-cracking of the naphtha in the second riser exerts on the converter as a whole.

This invention, which considers a single converter, has the drawback of the use of a single catalytic system for meeting opposing objectives. As for the batch riser, it would be desirable to use a lower activity catalytic system to maximize LCO, for the naphtha/DO riser it would be desirable to use a higher activity catalyst, suitable both for maximizing non-specified naphtha cracking, so as to obtain the necessary quality and recovering LPG production, and for cracking the recycled DO of the first fractionator, a stream reasonably refractory to the cracking. The differentiated catalytic system would also be necessary to enable the cracking of naphtha into high value, light olefins used as inputs for the petrochemical industry. This fact—the existence of a single converter—prevents the ideal optimization of the catalytic system, which limits the results that could be achieved through the configuration proposed in the patent application PI 0504321-2.

Another point is that the re-cracking of the naphtha and the DO occurs in a single riser. As these streams have a different nature, the ideal conditions for cracking them are also different. The fact of having a single riser prevents the optimization of the ideal catalytic cracking conditions of these streams, also limiting the results that could be achieved.

In addition, the converter operating reliability is reduced due to problems related to coke deposition on the equipment in the reaction and rectification section associated with lower riser temperature. In practice, the application of this process can be restricted to the processing of lighter feedstocks.

U.S. Pat. No. 7,632,977, which seeks simultaneously to increase the production of diesel and LPG in a single converter equipped with two risers (a low severity one for maximizing LCO and a high severity one for re-cracking naphtha), has the same disadvantage of using a single catalytic system for the feedstock and naphtha cracking. Furthermore, the DO recycling is sent to the low severity riser at a point above the feedstock injection, complicating the cracking of this stream, which requires greater severity for the converter. Therefore, do not expect a sharp reduction in the production of DO, which continues to be a problem for the refiner desiring to increase production of LCO diesel without the offsetting increase in the generation of DO. In addition, coke deposition problems in the low severity reaction section are compounded by the recycling of decanted oil, the vaporization of which is hindered by selected injection point.

Patent application PI 0605009-3 discloses a converter with a dual riser, with the difference that neither stream is recycled. The risers operate with different severities, and, therefore, the production profile and the qualities of the products are different between the two risers. The advantage of this design is the increased feedstock processing capacity of the unit, which represents an increased conversion supply to the refiner, since the second riser is not captive to the reprocessing streams of the first riser. However, the disadvantage is due to the production of a greater amount of non-specified naphtha and DO in the first riser.

Patent application U.S. 2006/0231458 discloses an FCC process that claims the use of a low contact time in the riser, focused on increasing the production of gasoline and middle distillates. This process uses a single riser and the contact time is from 1 to 5 seconds (however, in general, the low the contact time is characterized by times less than 2 seconds). This document cites some of the advantages of the low contact time, such as reduced non-selective reactions and a reduction of undesirable hydrogen transfer reactions leading to the formation of aromatics in the cracking products. The reduced production of fuel oil is achieved by selective DO recycling, that is, of a DO stream in which the triaromtic compounds (hydrocarbons with more than three aromatic rings) have previously been removed. This less aromatic DO stream will most certainly be crackable and will have a higher conversion than the raw DO from the fractionating tower, even with the low severity riser. However, this process requires a separation step (the patent application cites as possible separation methods: solvent extraction, chromatography and separation by means of membranes) to fractionate between “crackable” DO (saturates and mono/diaromatics) and the aromatic residue (tri-aromatics). Moreover, it is extremely difficult to obtain a quality LCO and naphtha specified as gasoline from cracking the feedstock in a single riser.

Document WO 01/60951 discloses an FCC process with two risers, the second riser being exclusively for reprocessing DO derived both from the first and second riser itself. As described in the patent application PI 0504321-2, the FCC process of WO 01/60951 reveals some disadvantages, such as:

    • the single recycled stream from the first riser, which operates at less severity, is the DO. As a result, the quality of the naphtha coming from this riser will not be good;
    • the naphtha and LCO streams of the first riser are not separated in the fractionator associated with this riser, but instead mixed with the same cuts obtained in the second DO riser. As a result, the LCO and naphtha obtained in the first riser are mixed with those obtained in the fractionator of the second riser. If the severity in both risers is low, the resultant LCO is of good quality, but reprocessing of the DO loses its effect, and the process usually has the drawback of high production of DO (and low quality naphtha). On the other hand, if the second riser is more severe, the DO conversion increases, but the quality of LCO of the second riser declines, and this stream is mixed with the LCO obtained in the first riser, with the overall result of a less significant improvement in the quality of the LCO;
    • in the design shown in the patent, the catalyst system of the two risers is the same;
    • coke deposition problems associated with the first riser can limit the reduction of the reaction temperature within it and also limit the quality of the feedstock to be processed.

U.S. Pat. No. 6,416,656 describes a process that uses a riser to increase the overall yield of diesel and LPG. In this process, the gasoline is re-cracked to increase the LPG yield, being injected at a point below the loading nozzle, the location being subject to greater severity in the riser, both in the reaction temperature and the ratio of the catalyst mass unit to the feedstock mass unit (C/O ratio). The process feedstock is injected at multiple points along the riser, reducing the contact time and thus increasing the LCO yield; moreover, in the upper part of the riser the temperature is already lower, as is the C/O ratio (due to the injection of additional feedstock into the riser), which also favors the yield and quality of the LCO. In addition, “quench” is injected into the top of the riser, for the purpose of stopping secondary reactions. The lower severity in the area for cracking the feedstock can lead to an unwanted increase in the production of DO; however, the patent indicates the possibility of recycling the DO together with the feedstock. The disadvantages of this technique are:

    • the use of a single riser for cracking the feedstock and recycling the naphtha and DO decreases the refiner's flexibility, as the operating window of the unit is shorter;
    • a considerable part of the installed riser is used for reprocessing, reducing the refinery's space for converting residues;
    • with a riser, although the injections take place at optimized temperature and C/O ratio points for each stream, the process is restricted to a single catalytic system for meeting such diverse objectives (maximizing diesel and LPG).

U.S. Pat. No. 7,316,773 describes an FCC process that seeks some of the characteristics that, during the development of its FCC technique for middle distillates, the Applicant considered to be important to the objective of maximizing middle distillates without the offsetting effects of increased DO and a loss of gasoline quality:

    • use of independent risers for re-cracking the cracked naphtha and the DO in order to optimize the reaction conditions for cracking streams with such as distinct chemical natures;
    • possibility of using catalytic systems specific to each objective: a catalytic system for cracking the feedstock at low severity (objective of maximizing quality LCO), and another catalytic system for cracking of naphtha in specific conditions (objective of correcting the quality of the gasoline in the FCC process for middle distillates).

However, U.S. Pat. No. 7,316,773 does not include and does not examine the potential as to a key characteristic for maximizing the allocation of the LCO obtained in the FCC to the refinery's diesel pool; one of the Applicant's objectives, namely, the segregation of better quality LCO (derived from the reactions from cracking the feedstock in low severity) from poorer quality LCO (derived from the reactions from cracking the DO at high severity);

The failure to meet this requirement significantly compromises the potential to improve the LCO quality of the process. This occurs because, unlike the Applicant's development that led to the present technique, U.S. Pat. No. 7,316,773 in any of its conceptions (FIGS. 1, 2 and 3 or combination thereof) performs the re-cracking of the DO in the same converter where the cracking of the feedstock is done, using the same separator vessel, the same rectifier, and consequently the same fractionator. This leads to the following two disadvantages:

First disadvantage: the mixing between the reaction product of the feedstock and the reaction product of the DO results in the LCO of both feedstocks being mixed. This leads to the following possibilities:

(a) the cracking of the DO is done with high severity, and, in this case, the DO yield is lower (positive effect), while LCO generated is of poorest quality and will be mixed with the LCO produced in the feedstock riser, resulting in reduced quality for this LCO (negative effect); or

(b), the DO is cracked at reduced severity, and, in this case, the quality of the mixed LCO of the two feedstocks is not as adversely affected; however, the overall DO yield from the unit remains high, therefore unsatisfactory.

Although the cracking in the DO riser is done at reduced severity, the LCO obtained from this stream is, generally, of poorer quality than that obtained by the cracking of the feedstock, because the DO has a larger proportion of aromatics in its composition—typically the only possible allocation for the thus obtained LCO would be the fuel oil pool, as a diluent. This means that whatever the reaction condition in the DO riser, the mixing of the LCO obtained in the DO riser with the LCO obtained in the feedstock riser is not a good practice for the purpose of maximizing the LCO quality.

Second disadvantage: despite the fact that U.S. Pat. No. 7,316,773 puts forth the possibility of using different catalytic systems for the feedstock and for the DO, this characteristic is not possible when both the two risers discharge the two product streams from the cracking into a single gas-solid separation vessel (separator vessel), as described in U.S. Pat. No. 7,316,773, without contamination occurring due to the mixing of the two catalytic systems with conflicting objectives. This mixing would occur in the separator vessel itself, in the rectification system, in the regenerator, and in the catalyst transfer standpipes, which, over time, would lead to the complete mixing of the two catalytic system inventories, both losing their characteristics—unless the equipment had physical dividers separating them over their entire height and the standpipes were independent.

The possibility of physical partitioning is raised in U.S. Pat. No. 7,316,773, only for the regenerator (in its FIG. 2) where it is provided for separating the catalytic system for the naphtha from the catalytic system used jointly for DO and fresh feedstock. In practice, this partitioning is not economically attractive, it is technically difficult, and its execution is improbable. Therefore, two converters would be necessary, as proposed by the Applicant in this technique. Nevertheless, according to the design of U.S. Pat. No. 7,316,773, the transfer lines of both reaction products must be interconnected, leaving the previously indicated first disadvantage.

Thus, the technique still requires a fluid catalytic cracking process with mixed feedstock aiming the production of middle distillates, which optimizes to the maximum the quality of the LCO obtained in cracking the feedstock, this high-quality LCO stream being isolated from other LCO streams deriving from the reprocessing of cracked fractions that, due to the chemical nature and because they require more stringent reaction, generate a poorer quality LCO.

Furthermore, despite mentioning the use of low contact time in the batch riser (less than 1.5 seconds), U.S. Pat. No. 7,316,773 does not associate the use of this variable with the benefit of enabling iso-conversion, the use of a higher reaction temperature, as proposed by the Applicant in the present technique. Without this technique, the entire reaction/rectification section for cracking the feedstock and the re-cracking of the DO would be subject to problems of coke deposition and loss of operating reliability, which could limit the range of possible batches that could be processed in the unit subjected to the technique made known in U.S. Pat. No. 7,316,773.

This present technique constitutes a development of the Applicant in relation to prior patent application PI 0504321-2, in that it incorporates:

    • the use of a second converter exclusively for re-cracking the naphtha and DO, allowing the use of a specific high activity catalyst system for these streams and another low activity one for the fresh feedstock, this latter this with the primary objective of ensuring the obtaining of a high quality LCO;
    • use of independent risers for the naphtha and for the DO, enabling greater flexibility for adjusting the optimal operating conditions for cracking each of these streams;
    • use of low contact time between the catalyst and the feedstock as a method for reducing the reaction severity in the fresh feedstock riser (for producing high quality LCO), entailing the benefits resulting from operating at a higher reaction temperature, namely: increased rectifier efficiency, better feedstock evaporation at the base of the riser, and greater operating reliability of the unit to minimize the risk of coke deposition in the equipment of the reaction/rectification equipment. This combination of benefits even permits the processing of heavier feedstocks.

Thus, the invention provides a process utilizing two distinct converters operating together in a coordinated manner to maximize diesel through the independent cracking of streams of naphtha and DO for the production of specified gasoline and reduced production of fuel oil through conversion of it into lighter products, in addition to facilitating the specification as Aromatic Residue (RARO) of its remaining fraction resulting from the high aromaticity obtained through re-cracking, exhibiting a benefit greater than that obtained with only a single converter.

This technique can be extended to refineries with two FCC units with some modifications in the equipment and alignments between units, being necessary only to set one of the two converters to low severity conditions and the other to high severity conditions. The catalytic systems would be different and appropriate to each objective.

The high quality LCO stream would be produced in the first and the naphtha and OD streams of the first, plus the batch of the unit, would re-cracked in the second. Depending on the converter model, the naphtha could be re-cracked in an independent riser. The DO would be mixed with the batch. If equipped with a single riser in the second converter, alternatively, the naphtha could be re-cracked in the “lift” section, below the injection of the batch mixture with DO.

Whatever the application, the use of this technique also offers the refiner the desired flexibility for changes in the local demand for derivatives. The “maximum LCO mode” condition can be modified to the “maximum gasoline mode,” and vice versa, simply by raising the reaction temperature and changing the feedstock injection point from the upper section of the riser to the lower section of the riser. In these conditions, the second converter will be used to crack part of the feedstock flow of the first, since the increased thermal demand of the first will limit the feedstock processing in it. The use of the two converters will make it possible to process the same total feedstock flow volume, regardless of the operating mode selected.

A more restricted application relates to its use in a refinery equipped with only one converter and that opts not to build a second converter to handle the stream of non-specified naphtha and the excess decanted oil coming from the converter operating for maximum LCO. In this case, the refiner takes advantage of the benefits to reduce the reaction severity in the converter via limited time contact, maintaining the highest TRX, and has the option of recycling part of the naphtha to the base of the riser. However, it will remain difficult to cope with the greater production of decanted oil: if one opts for recycling part of the DO or not to reduce the reaction severity so much, so as not to increase the DO yield, the consequence will be a deterioration in the quality of the LCO and less incorporation of this stream into the Diesel pool, the initial objective of the operation of the converter for maximum middle distillates. For the refiner who already utilizes the low severity in its converter by reducing TRX (and is able to balance the loss of quality of the cracked naphtha and the surplus decanted oil in its refining scheme), the replacement of this route by that of reduced contact time constitutes an advance toward greater operational reliability and the expansion of the unit's operating window.

SUMMARY OF THE INVENTION

The present invention relates to an FCC process using as heavy hydrocarbons or mixed hydrocarbons as feedstock, comprising two distinct converters, operating in a coordinated manner. A first converter operates with a low contact time in the riser, from 0.2 to 1.5 sec. (preferably from 0.5 to 1.0 sec.) making a higher reaction temperature possible even at low severity, from 510° C. to 560° C. (preferably from 530° C. to 550° C.) and with a catalyst suitable to the maximization of LCO. A second converter possesses a high activity catalytic system suited to the cracking of naphtha and DO and uses a conventional contact time (greater than 1.5 sec.)

The reaction temperatures of each are set according to the most recommended range for maximizing the cracking of each of the streams: 530° C. to 560° C. for the DO riser; and 540° C. to 600° C. for the naphtha riser. The process seeks both to maximize the production of LCO and its quality, as it makes it possible to produce gasoline with the required quality and to reduce fuel oil production, eliminating the typical problems encountered by refiners operating for maximized LCO. Moreover, the use of the low contact time in the first converter yields a superior quality LCO even from the use of higher reaction temperatures, therefore taking advantage of all the benefits associated with the use of higher TRX. The final result translated into increased operating reliability and a favorable thermal balance, opening space for increasing the feedstock temperature and for processing residual feedstocks.

BRIEF DESCRIPTION OF THE FIGURES

FIG. 1 schematically illustrates a flowchart of the proposed invention in its preferred embodiment wherein the second converter is equipped with two risers for re-cracking regardless of the naphtha and the DO in each riser.

FIG. 1 illustrates both the design of a new FCC unit aiming the production of diesel at the refinery and also illustrates the application of this invention at a refinery that already has two FCC converters, one of which is equipped with two risers.

FIG. 2 schematically illustrates a flowchart of an alternative solution of the present invention for application in a refinery that already has two FCC converters, where both only have one riser and for the case of not opting for the design of a second riser for one of the converters. It should be noted that in this case, the condition for re-cracking the naphtha and DO (in the same riser) is not optimized; however, FIG. 2 reflects a decision for lower investment in the implementation of this proposal.

FIG. 3 illustrates a block diagram of the invention applied to an existing FCC unit.

 DETAILED DESCRIPTION OF THE INVENTION

The present invention relates to an FCC process for maximizing middle distillates using as feedstock a mixture of hydrocarbons, which proposes the use of two different converters, operating in a coordinated manner for maximizing diesel, producing specified gasoline and at the same time reducing the production fuel oil, with a greater resultant benefit than that obtained with only one converter.

FIG. 1 schematically illustrates the proposal of the invention in its preferred embodiment.

According to FIG. 1, converter “A” (1) comprises a riser “I” (2), a separator vessel (3), a rectifier (4), a spent catalyst standpipe (5), a regenerator (6) and a regenerated catalyst standpipe (7). In converter “A,” (1) a feedstock (50a) consisting of heavy gas oil or atmospheric residue (AR) or a mixture of these streams is injected into larger load disperser (8) in the upper part of low severity riser “I,” (2) and this feedstock encounters the hot regenerated catalyst coming from the regenerator (6) through the pipeline, standpipe, (7). The regenerated catalyst is initially entrained by a “lift” fluid (51), which is generally steam, to assist the ascending flow of the catalyst into the riser “I” (2) up to the point of introduction of the feedstock. The flow rate must be set so as to ensure a stable flow of the catalyst. Cracking reactions take place in riser “I” (2) at a temperature between 510° C. and 560° C. (preferably from 530° C. to 550° C.) with a catalyst used to maximize the LCO, and a low contact time is used between the feedstock and the catalyst in the range 0.2 to 1.5 sec., preferably from 0.5 to 1.0 sec.

If so desired, there is the possibility of introducing a feedstock into the lower feedstock disperser (9) at the base of riser “I,” (2) subjecting it to a contact time with the catalyst of about 1.5 to 2.0 seconds, allowing the refiner to pass between the “maximum LCO mode” and the “maximum Gasoline mode.”

The reaction products (52) of the converter “A” (1) are lighter hydrocarbons than the feedstock (50a), such as gasoline, LPG, fuel gas (FG) e LCO, plus decanted oil (DO). These are sent to a fractionating tower “A” (11), where the LCO stream (57), with a quality suitable for being incorporated into the pool of the feedstocks for hydrotreatment for Diesel, is removed. At the top of fractionating tower “A,” (11) a light hydrocarbon stream exits (53), which is condensated and subsequently sent to a top vessel (12). Top vessel (12) separates a fraction of non-specified naphtha (NNE) (55) and a fraction of wet gases (54), comprised primarily of components of the FG and LPG range. The NNE fraction (55) is sent to naphtha disperser (29) at the base of a riser “II” (21) that constitutes part of the second converter “B” (20). Alternatively, a fraction of the NNE (56) can be sent to disperser (10) at the base of riser “I” (2) of converter “A” (1), where in this case it would serve as a “lift” fluid (naphtha-lift) for the catalyst, which requires a greater amount of lift steam (51) to reach to the upper part of the riser without the feedstock injection (which when vaporized ends up contributing to the upward flow of the catalyst). That is, in the option for the use of naphtha-lift (56), the consumption of steam (51) for the “lift” of the catalyst in riser “I” (2) is reduced, also reducing the generation of acid water in the fractionating section of converter “A” (1). The naphtha-lift (56) must be adjusted in relation to the quality of LCO (57) separated in fractionating tower “A” (11), so as not to compromise its incorporation into the refinery Diesel pool.

A decanted oil (DO) stream (58) exits at the bottom of fractionating tower “A” (11), which is sent to disperser (30) at the base of a riser “III” (28) of converter “B” (20). Converter “B” (20) comprises a riser “II” (21), a separator vessel (22), a rectifier (23), a spent catalyst standpipe (24), a regenerator (25), a regenerated catalyst standpipe (26) interconnecting regenerator (25) with riser “II” (21), a regenerated catalyst standpipe (27) interconnecting regenerator (25) with riser “III” (28), and a riser “III” (28).

In high severity riser “II” (21), the NNE (55) encounters the hot regenerated catalyst coming from the regenerator (25) through the pipe, standpipe (26). The regenerated catalyst is initially entrained by a “lift” fluid (59), which is generally steam, to assist the ascending flow of the catalyst into riser “II” (21). Cracking reactions take place in riser “II” (21) at a temperature between 540° C. and 600° C. with a high activity catalyst.

In riser “III” (28) at intermediate severity, the DO (58) encounters the hot regenerated catalyst coming from regenerator (25) through the pipe, standpipe (27). The regenerated catalyst is initially entrained by a “lift” fluid (60), which is generally steam, to assist the ascending flow of the catalyst into riser “III” (28).

Cracking reactions take place in riser “III” (28) at a temperature between 530° C. and 560° C. with a high activity catalyst, the same used in riser “II” (21). In addition, primarily in the case of application of this process to a refinery that already has the two converters (both of which process feedstocks), converter “B” (20) is expected to inject a feedstock (50b), consisting of heavy gas oil or atmospheric residue (AR), or a mixture of these streams, with the recycling of DO (58) of converter “A” (1) into the disperser (30) of riser “III” (28). Feedstock (50b) of converter “B” (20) can be the same as that used in converter “A” (1), or even a segregated batch with quality different from feedstock (50a).

Alternatively, it can be provided with a combination of independent dispersers for the feedstock (50b) and for the DO (58). In the case of setting the operating mode for maximum Gasoline, converter “A” (1) begins to produce a previously specified naphtha and, therefore, it is not necessary to re-crack this stream in riser “II” (21) of converter “B” (20). In this case a disperser (35) is provided in riser “II” (21) of converter “B” (20) for cracking feedstock (50c), comprising heavy gas or atmospheric residue (AR), or a mixture of these streams, which can be the same as those used in converter “A” (1) and riser “III” (28) of converter “B” (20), or even a segregated feedstock with a quality different from the feedstocks (50a) and (50b).

Reaction products (61) of converter “B” (20) are sent to a fractionating tower “B” (31) from which a stream exits, from the intermediate region of the tower, of low quality LCO (65), which is to be used as the refinery's fuel oil diluent. At the top of fractionating tower, “B” (31), a light hydrocarbon stream (62) exits, which is condensated and subsequently sent to a top vessel (32), which generates a stream (63) comprised primarily of components in the FG and LPG range. A non-stabilized naphtha stream (64) exits at the bottom of vessel (32). Sending streams (63) and (64) to the gas recovery area promotes the separation of FG and LPG, as well as the stabilization of the naphtha, which can then be incorporated into the refinery's Gasoline pool.

At the bottom of fractionating tower “B” (31), DO (66) exits with a higher aromatic content in its composition after re-cracking with greater severity in riser “III” (28), which enables its specification as an aromatic residue (DO for RARO), a residual product with a high aromatic and polyaromatic hydrocarbon content, which has a greater market value than the fuel oil, the main industrial application of which is for producing carbon black.

If the design is for a new unit (where there are no independent gas recovery areas for each converter), the wet gas fraction (54) leaving top vessel (12) of the converter “A” (1) can join the fraction of wet gases (63) exiting top vessel (32) of converter “B” (20) and continue to a common gas recovery area.

Converter “A” (1) operates with low contact time in riser “I” (2), in the range of 0.2 to 1.5 sec. (preferably from 0.5 to 1.0 sec.) and uses an appropriate catalyst for maximizing the LCO for the purpose of minimizing the aromatic content in the LCO, thus increasing its quality. These reaction products (52) are sent to a fractionating tower “A” (11), where LCO stream (57), with a quality suitable for being incorporated into the pool of feedstocks for Diesel hydrotreatment, is removed. Non-specified naphtha (NNE) (55) and DO (58), both removed from fractionating tower “A” (11) are sent to converter “B” (20), which has two independent risers, riser “II” (21) and riser “III” (28), the first being for cracking the NNE (55) and the second for cracking the DO (58). In separator vessel (22), risers “II” (21) and “III” (28) are coupled to a quick separation system, as proposed in the Applicant's patent PI 0704443-7, incorporated here as a reference.

Converter “B” (20) has a high activity catalytic system, suitable for cracking the NNE (55) and the DO (58), and as it has two separate risers, it allows the reaction temperatures of each to be set according to the range most recommended for maximizing the cracking of each of the streams: 530° C. to 560° C. for riser “III” (28) for DO (58) and 540° C. to 600° C. for riser “II” (21) for NNE (55).

The use of different catalytic systems, added to the fact of the use of three independent risers for cracking the feedstock (50a) and the NNE (55) and DO (58) streams produced in the converter “A” (1), provides ideal conditions for optimizing the desired results, which are: maximum production of LCO for diesel pool (57), minimum production of fuel oil (65), which is the LCO used as a diluent because it lacks the quality for the diesel pool and production of naphtha with a specification appropriate to gasoline pool (64). Moreover, the use of low contact time in converter “A” (1) makes it possible to obtain an LCO stream (57) with adequate quality for incorporation into the diesel pool based on the use of higher reaction temperatures, therefore, taking advantage of all the benefits associated with the use of the higher TRX. The final result translated into increased operating reliability and a favorable thermal balance, opening space for increasing the batch temperature and for processing residual feedstocks.

This process, as described in FIG. 1, can be used in new units, but is also perfectly suited to the case of improvements to an existing unit, where there is already an FCC unit in operation at the refinery. The converter can be put into low severity conditions for cracking the heavy vacuum gas oil (HVGO) feedstock (50a), as in the example of converter “A” (1). The NNE (55) and DO (58) removed from fractionating tower “A” (11) would be sent to the new converter “B” (20), equipped with two risers. A new fractionating tower “B” (31) would be provided to receive reaction products (62) from converter “B” (20). The existing gas recovery area, with minor modifications, would be suited for receiving the fuel gas (FG) and liquefied petroleum gas (LPG) produced in both converters. Another possible application of the process as described in FIG. 1 is for those refiners who already have two independent FCC units installed in their industrial park, at least one of these units being equipped with two risers. In this case, it is sufficient that the converter with a single riser is put in low severity conditions, as in example of converter “A” (1), and that interconnections are established for the NNE (55) and DO (58), removed from fractionating tower “A” (11), to the second converter with two existing risers, as in the example of converter “B” (20). With the interconnections between the wet gas streams of converter “A” at low severity and of converter “B” at high severity, (54) and (63) respectively, it would be possible for the two existing cold areas to accommodate the FG and LPG produced in both converters.

While the embodiment shown in FIG. 1 is the preferred one of this process, because it permits the benefit of the independent re-cracking of naphtha and DO generated in the low severity converter, a possible application of the invention is to a refinery having two FCC units, where both have only one riser. In this case and when not opting for improvements to an existing unit, specially when encompassing the design of a new riser for the converters, the scheme proposed in FIG. 2 is an alternative for the refinery to adopt the FCC route for diesel.

The process descriptions for converter “A” (1) and for fractionating tower “A” (11) in FIG. 2 are the same made in the scheme of FIG. 1.

In FIG. 2, however, the NNE (55) and DO (58) are re-cracked in the same riser (41) of converter “B” (40). In this case, the naphtha (55) is injected into disperser (47) at the base of riser (41), close to the arrival point of regenerated catalyst coming through the standpipe pipe (46) of regenerator (45). In the case of a single riser for the second converter, this is the ideal location for the naphtha, as it is the point of greatest severity of the riser, since the temperature of the mixture (catalyst+naphtha) is higher and the naphtha is subjected to a higher catalyst/naphtha at the base, as it is upstream of the DO and batch injections, which increases the hydrocarbon flow rate, and, therefore, reduces the effective catalyst/naphtha ratio for the same catalyst circulation. The DO (58) is injected into dispersion nozzles (48) located above the NNE (55) injection. In addition, feedstock (50b), which the refinery processed in this converter “B” (40), can continue to be cracked by injecting it together with the DO (58) into disperser (48). Alternatively, it can be provided with a combination of independent dispersers for the feedstock (50b) and for the DO (58).

Converter “B” (40) comprises a riser (41), a separator vessel (42), a rectifier (43), a spent catalyst standpipe (44), a regenerator (45) and a regenerated catalyst standpipe (46).

Reaction products (68) of converter “B” (40) are sent to a fractionating tower “B” (33) from which a stream exits, from the intermediate area of the tower, of low quality LCO (72), which is then to be used as the refinery's fuel oil diluent. At the top of fractionating tower “B” (33) a light hydrocarbon stream (69) exits, which is condensated and subsequently sent to a top vessel (34), which generates a stream (70) comprised primarily of components in the FG and LPG range. A non-stabilized naphtha stream (71) exits at the bottom of the vessel. The sending of streams (70) and (71) to the gas recovery area promotes the separation of FG and LPG, as well as the stabilization of the naphtha, which can then be incorporated into the refinery's Gasoline pool.

At the bottom of fractionating tower “B,” (33) DO (73) exits with a higher aromatic content in its composition after re-cracking with greater severity in riser (41), which enables its specification as an aromatic residue (DO for RARO), a residual product with a high aromatic and polyaromatic hydrocarbon content, which has a greater market value than the fuel oil, the main industrial application of which is for production of carbon black.

The wet gases of converter “A” (1) and converter “B” (40), (54) and (70) respectively, are subsequently interconnected so that the two existing cold areas can process the total wet gas flow without need for major modifications.

Converter “B” (40) has a high activity catalyst system, suitable for cracking NNE (55) and DO (58), but as it does not have two separate risers; it lacks the flexibility for independently seeing the reaction temperature for each stream. Riser (41) must operate at a temperature of 530° C. to 570° C. to facilitate the conversion of the DO, knowing that the naphtha will be subjected to a greater severity due to being injected at the base of riser (41).

In FIGS. 1 and 2, the flow of the catalyst/oil reactive mix is shown as ascending (in the risers). However, the Applicant has its catalyst/oil reactive mix descending flow technology, under Patent WO 2005/080531 (fully incorporated herein as reference), which can alternatively be used as a process and apparatus for achieving the above-described cracking reactions.

Not shown in FIGS. 1 and 2 is the injection of “quench” with a cooling fluid, such as water, at a point above the batch injections in any of the risers shown is an alternative that can be utilized as described in Patent PI 0504854-0 for the present process. Another design not explained by FIGS. 1 and 2 is the possibility of making additional cuts in the heavy naphtha (HN) and heavy cycle oil (HCO) fractionating towers, introducing greater adjustment flexibility between the quality of the cuts of the fractionator and that of the end derivatives of the refinery.

The block diagram, according to FIG. 3, illustrates the application of this innovation to a refinery with an existing FCC. The scheme illustrates the innovation, with the input of the feedstock (50a), such as heavy vacuum gas oil (HVGO), into an existing converter “A” (1). At the output of converter “A” (1) reaction products (52) are generated that continue to an existing fractionating tower “A” (11). Reaction products (52) of converter “A” (1) are lighter hydrocarbons than the feedstock (50a), such as gasoline, LPG, fuel gas (FG) and LCO, plus decanted oil (DO). These are sent to a fractionating tower “A” (11), where LCO stream (57), with a quality suitable for being incorporated into the diesel pool, is removed.

Non-specified naphtha streams (NNE) (55) and another DO one (58), as well as a wet gas stream (54), comprised primarily of FG and LPG are also separated by fractionating tower “A” (11). The NNE (55) and DO (58) streams continue to new converter “B” (20), equipped with two risers. After passing through converter “B” (20), reaction products (61) continue to a new fractionating tower “B” (31), where the wet gas fraction (63) comprised primarily of FG and LPG joins wet gases (54) of converter “A” (1) and continues to an existing gas recovery area (36). A second fraction exiting fractionating tower “B” (31) is comprised of a non-stabilized naphtha stream (64), which also continues to the gas recovery section (36), where it participates in the absorption of moisture and, following stabilization, continues to the refinery's Gasoline pool. Two more fractions are produced from this fractionating tower “B” (31), which are the low quality and high aromatic content LCO for diluent (65), and a fraction of DO for the aromatic residue (DO for the RARO) (66)

After passing through gas recovery section (36) FG (74), LPG (75) and Gasoline (76) fractions are produced.

The following example illustrates the operation and the production profile of this process, which must not be understood to be a restriction on other operating conditions.

Example 1

In this example, converter “A” (1) operates with a HGO (Heavy Gas Oil) feedstock (50a) at a reaction temperature of 520° C. and contact time of 0.5 sec, providing the following yield profile:

Riser I (2) @ t = 0.5 s and Yield profile, % p/p TRX = 520° C. Conversion (100% - 47.5 LCO - DO) FG 2.0 LPG 9.5 NNE 31.0 LCO 18.0 DO 34.5 Coke 5.0

In converter “B” (20), naphtha (55) and DO (58) coming from converter “A” (1) are re-cracked, respectively, in riser “II” (21) at 580° C. and in riser “III” (28) at 535° C. The following yield profile is obtained:

Yield profile in converter B (20), % w/w Naphtha Riser II (21) @ DO Riser III (28) @ 580° C. 535° C. Conversion — 64.0 FG 2.0 4.0 LPG 16.5 16.5 Gasoline 74.0 35.0 LCO for diluent 4.0 17.5 DO for RARO — 18.5 Coke 3.5 8.5

These data make it possible to calculate the unit's overall yield. The below table compares the results of this process with the typical FCC yield profile.

Overall Yield Profile in relation to GOP, % w/w FCC of Middle Conventional FCC Distillates Conversion 70.5 68.0 FG 5.0 4.0 LPG 19.0 20.0 Gasoline 41.0 35.0 LCO for diluent 8.5 7.5 LCO for “Diesel” 8.0 18.0 DO 13.0 — DO for RARO — 6.5 ΣCoke in the FCCs 5.5 9.0

As may be observed, the present case, even operating for the purpose of middle distillates, can recover FG and LPG yields (which in the case of a improvement in FCC units enables the use of the same existing gas recovery area). The gasoline yield, as desired, is reduced by about 15%, contributing to reducing the excess of this fuel in the market (which in turn can be easily reversed by adjusting the FCC operating mode to maximum gasoline, with the injection of the feedstock (50a) into lower disperser (9) of converter “A” (1) in case of a seasonal increase in gasoline demand). The contribution of the FCC unit to the diesel pool increases significantly, while at the same time the production of fuel oil is minimized through the marketing of the DO produced as aromatic residue (RARO).

The following example illustrates the benefit from use of the low contact time in the riser as the method for reducing the reactive severity in the converter geared towards producing LCO for diesel.

Example 2

This example describes three operating scenarios of converter “A” (1) obtained from experimental runs implemented on the Applicant's prototype FCC unit.

Run “A” represents a base operating case, with the use of a high activity catalyst and TRX of 530° C. Runs “B” and “C” represent the different routes for reducing reaction severity from the base case. In both, the reduced severity occurs both due to the use of a low activity catalyst and also due to an adjustment in the operational condition; however, they differ in the process variable used for the adjustment: run “B” shows the TRX reduction route, while the run “C” shows the contact time reduction route.

The main results are shown in the below table:

RUN A B C Case Base case Reduced Reduced TRX contact time Catalyst High activity Low activity Low activity TRX (° C.) 530 490 530 Contact time in the riser(s) 2.0 2.0 0.5 Feedstock temperature (° C.) 220 140 340 C/O 7.5 6.0 5.0 Dense phase temperature 695 695 695 (° C.) Yield FG 4.0 2.0 2.0 profile LPG 19.0 10.0 10.0 (% w) LNG 40.5 27.0 28.5 LCO 16.5 18.5 18.0 DO 13.0 36.5 37.0 Coke 7.0 6.0 4.5 Conversion (% w) 70.5 45 45 Delta-coke 0.9 1.0 0.9 Cetane of the LCO 24 33 33

It is apparent that both routes (“B” and “C”) were successful in reducing the conversion and consequently the production of a better-quality LCO compared to the base case.

All the runs are referenced on the same dense phase temperature, based on adjustment of the feedstock temperature in each run. It becomes possible to ascertain that, compared to the base case, in run “B” (TRX reduction route) there was a trend towards increased dense phase temperature, which had to be corrected by reducing the feedstock temperature from 80° C. Then, in run “C” (the contact time reduction route) there was, conversely, a downward trend in the dense phase temperature of the Regenerator that could be utilized to raise the feedstock temperature to 120° C.

Directly noting the resultant coke yield of the heat balance and of the feedstock temperature adjustment in each scenario, one finds that even in the TRX reduction route there is a 15% gap in the coke production relative to the base case. This gap is attributable to the replacement of the high activity catalytic inventory (greater tendency to form coke, or greater delta-coke) due to a low activity (lower delta coke) suitable for low conversion operation. The use of a low activity catalyst for the heat balance serves to cool the dense phase temperature; however, for run “B,” much of this benefit is consumed by the dense phase heating tendency of the Regenerator due to low TRX in the riser.

In run “C”, the coke production gap is 35% relative to the base case, significantly higher than that obtained by run “B.” This occurs because the low contact time route enables full utilization of the entire cooling of the dense phase temperature achieved by replacing the catalyst in the unit.

The benefit of using the low contact time is most evident when compared to the results of runs “B” and “C” (the two with the same catalyst). Both runs attain the same conversion and production level, and LCO quality. In run “C,” however, there is a thermal gap sufficient to raise the feedstock temperature for the same dense phase temperature to 200° C. This enabled a 25% lower coke production than that of run “B,” despite the higher TRX (which increases the coke demand of the converter) in run “C.” This exemplifies a way of harnessing thermal break that the low contact time route provides: the temperature rise of the feedtsock reduces the coke demand in the converter, with no use of part of the air of the combustion air blower. This exceeding air of the air blower could be used to increase the feedstock flow rate of the unit without having to invest in the Regeneration Section of the Converter. A comparison between runs “B” and “C” shows that the TRX reduction does not introduce the same gap and expansion of the operating window that reducing the contact time provides.

This example demonstrates (with runs “B” and “C” at the same conversion and the same LCO quality) the superiority, for the heat balance of the unit, of the use of the low contact time/higher TRX combination over the high contact time/lower TRX combination. Aside from enabling the increased feedstock temperature, the low contact time/higher TRX combination in an industrial unit also provides the aforementioned benefits: it increases operating reliability due to reduced coke deposition in the equipment of the reaction and rectification section, and enables the processing of residual feedstock in the unit.

Despite the contact time being a variable cited in the literature as one of the possible variables for adjusting the reaction severity, the approach of using of low contact time making it possible to work with a higher reaction temperature and thereby to gain all the benefits from the use of the low contact time/higher TRX combination even with the converter geared towards the production of middle distillates (when most middle distillate FCC processes described in the literature follow line of reduced TRX) is not found in the literature. This approach, learned by the Applicant over the development of its FCC process for maximizing diesel, is a central aspect of present invention, which will lead to a new riser design philosophy for FCC units encompassing cases of maximum LCO operation.

Thus, the contact time variable takes on a crucial importance in this process. Existing units can be optimized for the maximum LCO operation through the installation of feedstock injectors in the highest position of the riser. And in the case of new units, designed for campaigns set between maximum LCO or maximum gasoline, the use of injectors at different heights of the reactor will then be considered (offering flexibility to the unit), as it will be associated with all the benefits of using a higher TRX (favorable heat balance and operating reliability) even in a maximum LCO campaign.

In a context where the FCC is one of the main producers of gasoline in the refinery, and where the market for this derivative appears to be declining, while the diesel market is increasing and, moreover, the diesel quality specification is also increasing, the development of an FCC process that helps to offset the gasoline and fuel oil surplus in a market short of diesel has strategic importance and is essential to ensuring that refineries remain at full capacity. These are challenges that the present invention serves to meet.

Claims

1. FCC PROCESS FOR MAXIMIZING DIESEL, comprising a converter “A” (1) operating at low severity, aiming the production of best quality LCO to be incorporated into the refinery diesel pool, and converter “B” (20) operating at high severity, wherein converters “A” (1) and “B” (20) operate in a coordinated manner with the converter “B” (20) receiving non-specified naphtha (NNE) (55) and decanted oil (DO) (58) generated in converter “A” (1), for the purpose of adjusting the production and quality profile of the other products of the process, achieving in converter “B” (20) a reduction of the fuel oil production, gasoline specification and recovery of LPG from the production.

2. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the feedstock (50a) for converter “A” (1) comprises heavy gas oil from distillation, heavy gas oil from coke, or atmospheric residue or a mix of these streams.

3. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the feedstock (50b) for the second converter comprises heavy gas oil from distillation, heavy gas oil from coke, or atmospheric residue or a mix of these streams.

4. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the converter “A” (1) comprises a riser “I” (2), a separator vessel (3), a rectifier (4) and a regenerator (6), with the catalyst flowing through standpipe (5) of rectifier (4) to regenerator (6) and through standpipe (7) of regenerator (6) to riser “I” (2).

5. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the low severity in converter “A” (1) entails the use of a contact time between batch and catalyst in riser “I” (2) in a range between 0.2 and 1.5 sec., preferably between 0.5 and 1.0 sec.

6. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the converter “B” (20) comprises a riser “II” (21), a riser “III” (28), a separator vessel (22), a rectifier (23), and a regenerator (25), with the catalyst flowing through standpipe (24) of rectifier (23) to regenerator (25) and through standpipe (26) of regenerator (25) to riser “II” (21) and through standpipe (27) of regenerator (25) to riser “III” (28), this being the preferred embodiment for converter “B” (20).

7. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the feedstock (50c) to be injected into riser “II” (21) converter “B” (20), comprises heavy gas oil from distillation, heavy gas oil from coke, or atmospheric residue or a mix of these streams.

8. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the converter “B” (40), alternatively, comprises a single riser (41), a separator vessel (42), a rectifier (43), and a regenerator (45), with the catalyst flowing through standpipe (44) of separator vessel (42) to regenerator (45) and through standpipe (46) of regenerator (45) to riser (41).

9. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the converters “A” (1) and “B” (20) generate reaction products that continue to the different fractionating towers, to fractionating towers “A” (11) and “B” (31) respectively.

10. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the converters “A” (1) and “B” (40) generate reaction products that continue to fractionating towers “A” (11) and “B” (33) respectively.

11. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein in converter “A” (1), a feedstock (50a) is injected into disperser (8) in a position higher than low severity riser “I” (2) (in maximum LCO operation) and, optionally, is injected into a disperser (9) at the base of riser “I”, providing flexibility to the refiner to change to the maximum gasoline mode.

12. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the cracking reactions in riser “I” (2) take place at a temperature between 510° C. and 560° C., preferably between 530° C. and 550° C., using a low activity catalyst suitable for maximizing the LCO.

13. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the reaction products (52) of converter “A” (1) are lighter hydrocarbons than the feedstock (50a), such as gasoline, LPG, fuel gas (FG) and LCO, plus the decanted oil (DO), and these are sent to a fractionating tower “A” (11), from which LCO stream (57) with suitable quality is removed for incorporation into the diesel pool.

14. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein at the top of fractionating tower “A” (11), a wet gas stream (54) and a non-specified naphtha (NNE) (55) stream is generated, that is separated into a top vessel (12).

15. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the NNE fraction (55) is sent to the base of a, high severity, riser “II” (21), constituting part of second converter “B” (20); and alternatively an NNE fraction (56) is sent to the “lift” section of riser “I” (2), constituting part of first converter “A” (1).

16. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the NNE fraction (55) is sent to the base (area of high severity) of a riser (41), constituting part of second converter “B” (40); and alternatively an NNE fraction (56) is sent to the “lift” section of riser “I” (2), constituting part of first converter “A” (1).

17. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the DO stream (58) exits the bottom of a fractionating tower “A” (11) a DO stream (58) exits and is sent to the base of a riser “III” (28) of converter “B” (20).

18. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein a DO current (58) exits the bottom of a fractionating tower “A” (11), and is sent to a disperser (48) located above a disperser (47) for injecting NNE (55), into riser (41) of converter “B” (40).

19. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the reaction products (61) of converter “B” (20) are sent to a fractionating tower “B” (31) from which an LCO stream (65) for fuel oil diluent exits.

20. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the reaction products (68) of converter “B” (40) are sent to a fractionating tower “B” (33) from which an LCO stream (72) for fuel oil diluent exits.

21. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein a wet gas stream (63) and an non-stabilized naphtha stream with quality for gasoline (64) are generated at the top of fractionating tower “B” (31), and are separated in a top vessel (32).

22. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1 wherein a wet gas stream (70) and an non-stabilized naphtha stream with quality for gasoline (71) are generated at the top of fractionating tower “B” (33), and are separated in a top vessel (34).

23. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein DO (66) with specification for aromatic residue (DO for RARO) is generated at the top of fractionating tower “B” (31).

24. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein DO (73) with specification for aromatic residue (DO for RARO) is generated at the bottom of fractionating tower “B” (33).

25. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the fraction of wet gases (63) exiting from the top (32) joins a wet gas stream (54) from converter “A” (1) and continues to a common gas recovery area, or alternatively to two different gas recovery areas, if there are such.

26. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the fraction of wet gases (70) exiting from the top (34) joins a wet gas stream (54) from converter “A” (1) and continues to a common gas recovery area, or alternatively to two different gas recovery areas, if there are such.

27. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the non-specified naphtha (NNE) (55) and the decanted oil (DO) (58), both removed from fractionating tower “A” (11), will be sent for re-cracking in converter “B” (20), which has two independent risers, riser “II” (21) and riser “III” (28), the first being for cracking NNE (55) and the second for cracking DO (58).

28. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the non-specified naphtha (NNE) (55) and the decanted oil (DO) (58), both removed from fractionating tower “A” (11), will be sent for re-cracking in converter “B” (40), which has a single riser (41), in which NNE (55) is cracked in its “lift” section and DO (58) is cracked in the upper section.

29. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein converter “B” (20) has an alternative to processing fresh feedstock (50b) in the same disperser (30) used for DO (58), or in another disperser level in riser “III” (28).

30. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein converter “B” (40) has an alternative to processing fresh feedstock (50b) in the same disperser (48) used for DO (58), or in another disperser level in riser (41).

31. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein converter “B” (20) has an alternative to processing fresh feedstock (50c) in the same disperser (35) of riser “II” (21), if the operating mode of converter “A” (1) is changed to maximization of Gasoline.

32. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein converter “B” (20) has high activity catalytic system, suitable to cracking NNE (55) and DO (58) and different from the low activity catalytic system used in converter “A” (1) for cracking a feedstock (50a).

33. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein converter “B” (40) has high activity catalytic system, suitable to cracking NNE (55) and DO (58) and different from the low activity catalytic system used in converter “A” (1) for cracking the feedstock (50a).

34. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, comprising two independent risers in converter “B” (20), which allows the reaction temperatures of each to be adjusted according to the most recommended range to maximize the cracking of the NNE (55) and DO (58) streams in each one of them, this being the preferred embodiment for cracking these streams.

35. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein riser “II” (21) operates with a reaction temperature between 540° C. and 600° C.

36. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein riser “III” (28) operates with a reaction temperature between 530° C. and 560° C.

37. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, comprising it has a single riser (41) in converter “B” (40) that operates at a reaction temperature between 530° C. and 570° C., with the cracking of NNE (55) being done in the “lift” section at the base of riser (41) and the cracking of DO (58) being done at a higher point.

38. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein riser “III” (28) for re-cracking DO (58) operates with a high reaction temperature and contact time, meeting the objective of reducing DO production, but generating some amount of poorer quality LCO (65).

39. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein riser (41) for re-cracking NNE (55) and DO (58) operates with high reaction temperature and contact time, meeting the objective of reducing DO production, but generating some amount of poorer quality LCO (72).

40. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the poorer quality LCO (65) generated by re-cracking DO (58) in converter “B” (20) is isolated from the high quality LCO (57) produced in the cracking of the feedstock (50a) at low severity in converter “A” (1).

41. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein poorer quality LCO (72) generated by re-cracking DO (58) together with NNE (55) in riser (41) of converter “B” (40) is isolated from the high quality LCO (57) produced in the cracking of the feedstock (50a) at low severity in converter “A” (1).

42. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, comprising a gas recovery area (36) where fractions of fuel gas, (74), LPG (75) e Gasoline (76) are produced.

43. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, wherein the flow of the reactive catalyst/oil mixture in any of the converters is downwards.

44. FCC PROCESS FOR MAXIMIZING DIESEL, according to claim 1, which can be applied to an existing FCC unit.

Patent History
Publication number: 20140034551
Type: Application
Filed: Apr 15, 2011
Publication Date: Feb 6, 2014
Applicant: PETROLEO BRASILEIRO S.A. - PETROBRAS (Rio de Janeiro, RJ)
Inventors: Emanuel Freire Sandes (Icarai Niteroi), Leandro Morais Silva (Rio de Janeiro), William Richard Gilbert (Rio de Janeiro), Jose Geraldo Furtado Ramos (Rio de Janeiro)
Application Number: 14/111,145
Classifications
Current U.S. Class: Catalyst In Multiple Stages (208/74)
International Classification: C10G 11/18 (20060101);