Patent Application
20080110801
The field of this invention is heating a hydrocarbon stream entering a reaction zone.
Hydrocarbon conversion processes that are exothermic or endothermic can be employed in the petroleum refining or petrochemical production industry. An exemplary hydrocarbon conversion process for improving the octane quality of hydrocarbon feedstocks is catalytic reforming where the primary product of reforming being motor gasoline or a source of aromatics for petrochemicals. The art of catalytic reforming is well known and a brief detailed description is provided below.
Generally, in catalytic reforming, a feedstock is admixed with a recycle stream comprising hydrogen to form what is commonly referred to as a combined feed stream, and the combined feed stream is contacted with a catalyst in a reaction zone. The usual feedstock for catalytic reforming is a petroleum fraction known as naphtha and having an initial boiling point of about 82° C. (about 180° F.) and an end boiling point of about 203° C. (about 400° F.). The catalytic reforming process is particularly applicable to the treatment of straight run naphthas comprised of relatively large concentrations of naphthenic and substantially straight chain paraffinic hydrocarbons, which are subject to aromatization through dehydrogenation and/or cyclicization reactions.
Reforming may be defined as the total effect produced by dehydrogenation of cyclohexanes and dehydroisomerization of alkylcyclopentanes to yield aromatics, dehydrogenation of paraffins to yield olefins, dehydrocyclicization of paraffins and olefins to yield aromatics, isomerization of n-paraffins, isomerization of alkylcycloparaffins to yield cyclohexanes, isomerization of substituted aromatics, and hydrocracking of paraffins. Further information on reforming processes may be found in, for example, U.S. Pat. Nos. 4,119,526 (Peters et al.); 4,409,095 (Peters); and 4,440,626 (Winter et al.).
A catalytic reforming reaction is normally effected in the presence of catalyst particles including one or more Group VIII (IUPAC 8-10) noble metals (e.g., platinum, iridium, rhodium, palladium) and a halogen combined with a porous carrier, such as a refractory inorganic oxide.
In a common form, the reforming process can employ the catalyst particles in several reaction zones interconnected in a series flow arrangement. There may be any number of reaction zones, but usually the number of reaction zones is 3, 4 or 5. Because reforming reactions generally occur at an elevated temperature, and are generally endothermic, each reaction zone usually has associated with it one or more heating zones, which heat the reactants to the desired reaction temperature. As a consequence of these considerations, a common process flow through the train of heating and reaction zones in a 3-reactor catalytic reforming process is as follows.
A naphtha-containing feedstock can admix with a hydrogen-containing recycle gas to form a combined feed stream, which may pass through a combined feed heat exchanger. In the combined feed heat exchanger, the combined feed can be heated by exchanging heat with the effluent of the third reactor. The heating of the combined feed stream that occurs in the combined feed heat exchanger is generally, however, insufficient to heat the combined feed stream to the desired inlet temperature of the first reactor. Consequently, after exiting the combined feed heat exchanger and prior to entering the first reactor, the combined feed stream often requires additional heating. This additional heating occurs in a heater, which is commonly referred to as a charge heater, which can heat the combined feed stream to the desired inlet temperature of the first reactor.
The combined feed stream may then pass to and through the first reactor. Because of the endothermic reforming reactions that occur in the first reactor, generally the temperature of the effluent of the first reactor falls not only to less than the temperature of the combined feed to the first reactor, but also and more importantly, to less than the desired inlet temperature of the second reactor. Therefore, the effluent of the first reactor can pass through another heater, which is commonly referred to as the first interheater, and which can heat the first reactor effluent to the desired inlet temperature of the second reactor.
On exiting the first interheater, the first reactor effluent usually enters the second reactor. As in the first reactor, endothermic reactions cause another decline in temperature across the second reactor. Generally, however, the temperature decline across the second reactor is less than the temperature decline across the first reactor, because the reactions that occur in the second reactor are generally less endothermic than the reactions that occur in the first reactor. Despite the somewhat lower temperature decline across the second reactor, the effluent of the second reactor is nevertheless still at a temperature that is less than the desired inlet temperature of the third reactor.
Consequently, the effluent of the second reactor can pass through another heater, which is commonly referred to as the second interheater, and then may pass to the third reactor.
In the third reactor, endothermic reactions cause yet another temperature decline, which is generally less than that across the second reactor, for the like reason that the temperature decline across the second reactor is generally less than that across the first reactor. The effluent of the third reactor can pass to the previously mentioned combined feed exchanger, where the effluent of the third reactor may be cooled by exchanging heat with the combined feed stream.
Generally, it is also known that a reforming unit can operate with different feed inlet temperatures for each of the reactors. Typically, such a unit has a train of three, four or five pairs of heaters and reactors that contain beds of catalyst, preferably fixed or moving beds, but many of the various possible combinations of different inlet temperatures, which together form what is usually called the temperature profile of the unit, are perhaps best illustrated with a three-reactor unit. If the inlet temperatures of all three reactors are the same, then the temperature profile is commonly called flat. Otherwise, the reactors can be operated with a non-flat or skewed reactor inlet temperature profile. As an example, if the inlet temperature of the first reactor is less than the inlet temperature of the second reactor, which is in turn less than the inlet temperature of the third reactor, then the profile of the reactor inlet temperatures is usually said to be ascending. If the first inlet temperature is more than the second inlet temperature, which is more than the third inlet temperature, then the profile is normally called descending. If the second inlet temperature is more than both the first and third inlet temperatures, then the profile can be said to resemble a hill. If the second inlet temperature is less than both the first and third inlet temperatures, then the profile may be said to look like a valley.
The most common reason for operating with a non-flat (i.e., skewed) reactor inlet temperature profile is to allocate the required heat duty among the heaters in the heater-reactor train. Ideally, all of the heaters are individually delivering heat at approximately the same percentage of their individual design duties. When each heater is operating at the same percentage of its design duty, as any other heater in the train is operating as a percentage of that other heater's design duty, then the heater duties are said to be “balanced.” Of course, a heater should not, as a general rule, be operated in excess of its design duty, that is the percentage should generally be less than or equal to 100%. A flat profile could result in imbalance of the operating duties of the heaters in the train, if some of the operating variables such as feedstock quality or throughput differ significantly from their design values, or if flow maldistribution or mechanical problems cause the performance of a reactor to fall significantly below its expected performance.
An illustration of attempting to balance heater duties in a commercial continuous reforming process by skewing reactor inlet temperatures is described in the article by Richard Lee, et al. entitled “Reforming Processes, Maximizing Profitability,” which begins at page 151 in Volume 47 of the Encyclopedia of Chemical Processing and Design, edited by John J. McKetta and published by Marcel Dekker, Inc., New York in 1994. In the example in the article by Lee et al., a valley-shaped profile of reactor inlet temperatures is recommended, where the inlet temperatures of parallel reactors 1 and 2 are the same and greater than the inlet temperature of reactor 3, which is less than the inlet temperature of reactor 4. Reactor 4's inlet temperature may be the same as or less than that of reactors 1 and 2. The largest difference between the reactors' 1, 2, and 4 inlet temperatures and the reactor 3 inlet temperature is 14° C. (26° F.). The Lee et al. article also teaches that the magnitude of the differences between the gasoline range product (C5+yield) when running an equal (that is, flat) reactor inlet temperature profile versus a staggered (that is, skewed) reactor inlet temperature profile is expected to be no more than 0.5% of feed.
A reforming process furnace may include multiple cells in which the feed to the reactors can be heated in the radiant section of the cell while steam is typically generated in the convection section of the heater. Generally, the heater capital cost is typically more than 20% of the unit cost. Because the feed may be heated in the radiant section of the heater, a significant quantity of the fired fuel, e.g., 30% of the fuel, may be actually used to generate steam instead of heating the process feed. Desirably, directing more of the heat into the process can reduce the costs of the heater and fuel.
U.S. Pat. No. 6,106,696 by David Fecteau and Kenneth Peters discloses the possible elimination of a heater in a reforming unit. The feed is heated and vaporized in the combined feed exchanger and charged to the first reactor directly at a relatively low temperature. The combined feed exchanger outlet temperature can be less than 482° C. (900° F.). The relatively low temperature at the combined feed exchanger outlet can lead to a higher reaction temperature requirement for the subsequent reactors if overall catalyst loading is constant and the first heater duty is higher and unbalanced as compared to the other heaters.
However, generally it would be desirable to increase the feed temperature to higher than the combined feed exchanger outlet to keep the subsequent reactor inlet temperature lower without increasing the catalyst loading. It can also be desirable to skew the reactor inlet temperature to balance the heating duty requirement for each heater cell.
One exemplary process can include passing a hydrocarbon stream through a reforming unit. The reforming unit may include a heater, which in turn includes a convection section and a radiant section, and a plurality of reforming reaction zones. Generally, the hydrocarbon stream is heated in the convection section for reacting in one of the reforming reaction zones to which the hydrocarbon stream is sent and the hydrocarbon stream is heated in the radiant section of the heater for reacting in the other reforming reaction zone to which the hydrocarbon stream is sent.
Another exemplary reforming process may include sending a stream including hydrocarbons though a reforming unit. The reforming unit may include at least one heater and a plurality of reforming reaction zones. Generally, the at least one heater includes a convection section and a radiant section where at least about 90% of heat transferred from the at least one heater to the hydrocarbon stream entering one of the reforming reaction zones is from one or more convection sections of the at least one heater.
An exemplary refinery or petrochemical production facility may include a reforming unit. Generally, the reforming unit includes at least one heater including a convection section and a radiant section. The convection section may include at least one convection tube having an inlet and an outlet, and the radiant section including a burner and at least one radiant tube having an inlet and an outlet. The reforming unit may further include a plurality of reforming reaction zones in a series wherein each reaction zone has an inlet and an outlet. Generally, the first reaction zone inlet is for receiving a hydrocarbon stream from the outlet of the convection tube, and a second reaction zone inlet is for receiving the hydrocarbon stream from the outlet of the radiant tube.
Yet another exemplary process can include passing a hydrocarbon stream through a reforming unit. The reforming unit may include a heater, which in turn can include a convection section and a radiant section, and a plurality of reforming reactors. Generally, the hydrocarbon stream is heated only in the convection section and not the radiant section of the heater before entering one of the reforming reactors.
A still further exemplary process can include operating a heater including a convection section and a radiant section, operating a plurality of reaction zones in series, passing a hydrocarbon stream through the at least one convection tube directly into the inlet of one of the zones, and passing the hydrocarbon stream through the at least one radiant tube directly into the inlet of the other zone. The convection section may include at least one convection tube having an inlet and an outlet and the radiant section may include at least one radiant tube having an inlet and an outlet. Moreover, each reaction zone may have an inlet.
Desirably, using the convection section of one or more heaters to heat the feed to one reactor and setting the inlet temperatures of the reactors, it is possible to replace a radiant section of a heater or furnace with a convection section. This can reduce the capital and catalyst cost, and fuel and flue gas flow at the same time, so the emissions from the unit (e.g., CO2, SOX, NOX) may also be reduced. Also, utilizing the convection sections and a skewed temperature profile can permit the shift of heat duties from the front end to the back end. Thus, the size of the furnaces can be standardized to reduce capital costs. Moreover, such an arrangement can obtain increases in yield, such as an increase of 0.1%. Moreover, the embodiments herein may allow a ratio of heater radiant sections to reaction zones of less than 1:1, such as 3:4 or 2:3.
As discussed above, U.S. Pat. No. 6,106,696 discloses a reforming process that employs at least two moving bed reaction zones, which preferably employs no heating between the combined feed exchanger and the lead reaction zone. U.S. Pat. No. 6,106,696 is hereby incorporated by reference in its entirety.
As used herein, the term “hydrocarbon stream” can be a stream including various hydrocarbon molecules, such as straight-chain, branched, or cyclic alkanes, alkenes, alkadienes, and alkynes, and optionally other substances, such as gases, e.g., hydrogen, or impurities, such as heavy metals. The hydrocarbon stream may be subject to reactions, e.g., reforming reactions, but still may be referred to as a hydrocarbon stream, as long as at least some hydrocarbons are present in the stream after the reaction. Thus, the hydrocarbon stream may include streams that are subjected to, e.g., a hydrocarbon stream effluent, or not subjected to, e.g., a naphtha feed, one or more reactions. As used herein, a hydrocarbon stream can also include a hydrocarbon feedstock, a feed, a feed stream, a combined feed stream or an effluent. Moreover, the hydrocarbon molecules may be abbreviated C1, C2, C3 . . . Cn where “n” represents the number of carbon atoms in the hydrocarbon molecule.
As used herein, the term “directly” can mean exiting a heater to a reaction zone without any substantial heat input from, e.g., the radiant or convection section of a heater or a heat exchanger.
As used herein, the term “radiant section” generally refers to a section of a heater receiving about 35- about 65% for fouled tubes or about 45- about 65% for relatively clean tubes of the heat, primarily by radiant and convective heat transfer, released by, e.g., the fuel gas burned by the heater.
As used herein, the term “convection section” generally refers to a section of a heater receiving about 10- about 45% of the heat, primarily by convective and radiant heat transfer by, e.g., the flue gas, released by the fuel gas burned by the heater. Typically, about 7- about 15% of the heat is lost through the stack, so usually no more than about 93% of the heat released by the fuel is utilized in the radiant and convection sections.
As used herein, the term “send” or “sent” with respect to a fluid can mean transferring a fluid from one location to another by means such as pumping or compressing, or by utilizing gravity.
As used herein, the term “heater” can include a furnace, a charge heater, or an interheater. A heater can include at least one burner and can include at least one radiant section, at least one convection section, or a combination of at least one radiant section and at least one convection section.
Generally, the embodiments disclosed herein are applicable to multiple reaction systems with multiple heaters, such as various hydrocarbon conversion processes, including those that are exothermic and endothermic. As an example, the embodiments of the invention could be applicable for an exothermic process with multiple reaction zones where an ascending temperature profile would be desirable. Preferably, the embodiments disclosed herein are applicable for endothermic reforming processes.
Generally, a hydrocarbon feedstock that is charged for a reforming process includes naphthenes and paraffins that boil within the gasoline range. The preferred charge stocks are naphthas consisting principally of naphthenes and paraffins, although, in many cases, aromatics also can be present. This preferred class includes straight-run gasolines, natural gasolines, synthetic gasolines, and the like. As an alternative embodiment, it is frequently advantageous to charge thermally or catalytically cracked gasolines or partially reformed naphthas. Mixtures of straight-run and cracked gasoline-range naphthas can also be used to advantage. The gasoline-range naphtha charge stock may be a full-boiling gasoline having an initial boiling point of about 40- about 82° C. (about 104- about 180° F.) and an end boiling point within the range of about 160- about 220° C. (about 320- about 428° F.), or may be a selected fraction thereof which generally can be a higher-boiling fraction commonly referred to as a heavy naphtha—for example, a naphtha boiling in the range of about 100- about 200° C. (about 212- about 392° F.). In some cases, it is also advantageous to charge pure hydrocarbons or mixtures of hydrocarbons that have been recovered from extraction units—for example, raffinates from aromatics extraction or straight-chain paraffins—which are to be converted to aromatics. In some other cases, the feedstock may also contain light hydrocarbons that have 1-5 carbon atoms, but since these light hydrocarbons cannot be readily reformed into aromatic hydrocarbons, these light hydrocarbons entering with the feedstock are generally minimized.
One exemplary feedstock that can be converted by these processes disclosed herein generally include a stream, which may be a naphtha, including, in percent by weight based on the total weight of hydrocarbons in the stream, components disclosed in Table 1:
Generally, the combined feed stream, or the hydrocarbon feedstock if no hydrogen is provided with the hydrocarbon feedstock, enters a heat exchanger at a temperature of generally about 65- about 177° C. (about 150- about 350° F.), and more usually about 93- about 121° C. (about 200- about 250° F.). Because hydrogen is usually provided with the hydrocarbon feedstock, this heat exchanger may be referred to herein as the combined feed heat exchanger, even if no hydrogen is supplied with the hydrocarbon feedstock. Generally, the combined feed heat exchanger heats the combined feed stream by transferring heat from the effluent stream of the last reforming reactor to the combined feed stream. Preferably, the combined feed heat exchanger is an indirect, rather than a direct, heat exchanger, in order to prevent valuable reformate product in the last reactor's effluent from intermixing with the combined feed, and thereby being recycled to the reforming reactors, where the reformate quality could be degraded.
Although the flow pattern of the combined feed stream and the last reactor effluent stream within the combined feed heat exchanger could be completely cocurrent, reversed, mixed, or cross-flow, the flow pattern is preferably countercurrent. By a countercurrent flow pattern, it is meant that the combined feed stream, while at its coldest temperature, contacts one end (i.e., the cold end) of the heat exchange surface of the combined feed heat exchanger while the last reactor effluent stream contacts the cold end of the heat exchange surface at its coldest temperature as well. Thus, the last reactor effluent stream, while at its coldest temperature within the heat exchanger, exchanges heat with the combined feed stream that is also at its coldest temperature within the heat exchanger. At another end (i.e., the hot end) of the combined feed heat exchanger surface, the last reactor effluent stream and the combined feed stream, both at their hottest temperatures within the heat exchanger, contact the hot end of the heat exchange surface and thereby exchange heat. Between the cold and hot ends of the heat exchange surface, the last reactor effluent stream and the combined feed stream flow in generally opposite directions, so that, in general, at any point along the heat transfer surface, the hotter the temperature of the last reactor effluent stream, the hotter is the temperature of combined feed stream with which the last reactor effluent stream exchanges heat. For further information on flow patterns in heat exchangers, see, for example, pages 10-24 to 10-31 in Perry's Chemical Engineers' Handbook, Sixth Edition, edited by Robert H. Perry et al., published by McGraw-Hill Book Company in New York, in 1984, and the references cited therein.
Generally, the combined feed heat exchanger operates with a hot end approach that is generally less than about 56° C. (about 100° F.), preferably less than about 33° C. (about 60° F.), and more preferably less than about 28° C. (about 50° F.). As used herein, the term “hot end approach” is defined as follows: based on a heat exchanger that exchanges heat between a hotter last reactor effluent stream and a colder combined feed stream, where T1 is the inlet temperature of the last reactor effluent stream, T2 is the outlet temperature of the last reactor effluent stream, t1 is the inlet temperature of the combined feed stream, and t2 is the outlet temperature of the combined feed stream. As used herein, for a countercurrent heat exchanger, the “hot end approach” is defined as the difference between T1 and t2. In general, the smaller the hot end approach, the greater is the degree to which the heat in the last reactor's effluent is exchanged to the combined feed stream.
Although shell-and-tube type heat exchangers may be used, another possibility is a plate type heat exchanger. Plate type exchangers are well known and commercially available in several different and distinct forms, such as spiral, plate and frame, brazed-plate fin, and plate fin-and-tube types. Plate type exchangers are described generally on pages 11-21 to 11-23 in Perry's Chemical Engineers' Handbook, Sixth Edition, edited by R. H. Perry et al., and published by McGraw Hill Book Company, in New York, in 1984.
In one embodiment, the combined feed stream can leave the combined feed heat exchanger at a temperature of about 399- about 516° C. (about 750- about 960° F.) to enter one or more convection sections of at least one heater, or a first heater. Generally, the feed stream enters the convection section at its top portion where the flue gases are at their coldest temperature and exits at the lower portion of the convection section where the flue gases are at their hottest temperature. Alternatively, the feed stream can enter the convection section at its lower portion where the flue gases are at their hottest temperature and exit at the higher portion of the convection section where the flue gases are at their coldest temperature. Alternatively still, the feed stream can enter and exit at the top or at the bottom of the convection section. The temperature of the combined feed stream leaving the convection section, which is also the inlet temperature of the first reaction zone, is generally about 482- about 549° C. (about 900- about 1020° F.), preferably about 518- about 538° C. (about 965- about 1000° F.).
One benefit of the present embodiment is the flexibility to not control the temperature at the convection section outlet. Rather, control of product quality can be obtained by adjusting the rest of the reaction zone inlet temperatures. In the case where independent temperature control of the first reaction zone is desired, control of the convection section process outlet temperature can be achieved by designing a combined feed exchanger with a hot-side or a cold-side bypass. A portion of the last reactor effluent stream or of the combined feed stream may bypass the combined feed exchanger. Alternatively, control can be obtained by a minor adjustment of excess air to the heater combined with a fine control by using a small hot side bypass on the combined feed heat exchanger.
This invention can be particularly applicable to the catalytic reforming of hydrocarbons in a reforming reaction system having at least two catalytic reaction zones where at least a portion of the reactant stream and at least a portion of the catalyst particles flow serially through the reaction zones. Reaction systems having multiple zones generally take one of two forms: a side-by-side form or a stacked form. In the side-by-side form, multiple and separate reaction vessels, each including a reaction zone, can be placed along side each other. In the stacked form, one common reaction vessel may contain the multiple and separate reaction zones that are placed on top of each other.
Although the reaction zones can include any number of arrangements for hydrocarbon flow such as downflow, upflow, and crossflow, the most common reaction zone to which this invention is applied can be radial flow. A radial flow reaction zone generally consists of cylindrical sections having varying nominal cross-sectional areas, vertically and coaxially disposed to form the reaction zone. Briefly, a radial flow reaction zone typically includes a cylindrical reaction vessel containing a cylindrical outer catalyst retaining screen and a cylindrical inner catalyst retaining screen that generally are both coaxially-disposed within the reaction vessel. The inner screen can have a nominal, internal cross-sectional area that is less than that of the outer screen, which may have a nominal, internal cross-sectional area that is less than that of the reaction vessel. The reactant stream can be introduced into the annular space between the inside wall of the reaction vessel and the outside surface of the outer screen. The reactant stream can pass through the outer screen, flow radially through the annular space between the outer screen and the inner screen, and pass through the inner screen. The stream that may be collected within the cylindrical space inside the inner screen can be withdrawn from the reaction vessel. Although the reaction vessel, the outer screen, and the inner screen may be cylindrical, they may also take any suitable shape, such as triangular, square, oblong, or diamond, depending on many design, fabrication, and technical considerations. For example, it is common for the outer screen to not be a continuous cylindrical screen but to instead be an arrangement of separate, elliptical, tubular screens called scallops that may be arrayed around the circumference of the inside wall of the reaction vessel. The inner screen is commonly a perforated center pipe that is covered around its outer circumference with a screen.
Illustrative reaction vessels that have stacked reaction zones and that may be used to practice this invention are shown in U.S. Pat. Nos. 3,706,536 (Greenwood, et al.) and 5,130,106 (Koves et al.), the teachings of which are incorporated herein by reference in their entirety. Transfer of the gravity-flowing catalyst particles from one reaction zone to another, the introduction of fresh or regenerated catalyst particles, and the withdrawal of coke-containing spent catalyst particles may be effected through catalyst transfer conduits.
Generally, the reforming reactions are normally effected in the presence of catalyst particles comprised of one or more Group VIII (IUPAC 8-10) noble metals (e.g., platinum, iridium, rhodium, and palladium) and a halogen combined with a porous carrier, such as a refractory inorganic oxide. U.S. Pat. No. 2,479,110 (Haensel), for example, teaches an alumina-platinum-halogen reforming catalyst. Although the catalyst may contain about 0.05- about 2.0 wt-% of Group VIII metal, a less expensive catalyst, such as a catalyst containing about 0.05- about 0.5 wt-% of Group VIII metal may be used. The preferred noble metal is platinum. In addition, the catalyst may contain indium and/or a lanthanide series metal such as cerium. The catalyst particles may also contain about 0.05- about 0.5 wt-% of one or more Group IVA (IUPAC 14) metals (e.g., tin, germanium, and lead), such as described in U.S. Pat. No. 4,929,333 (Moser et al.), U.S. Pat. No. 5,128,300 (Chao et al.), and the references cited therein. Generally, the halogen is normally chlorine and the alumina is commonly the carrier. Preferred alumina materials are gamma, eta, and theta alumina, with gamma and eta alumina generally being most preferred. One property related to the performance of the catalyst is the surface area of the carrier. Preferably, the carrier has a surface area of about 100- about 500 m2/g. The activity of catalysts having a surface area of less than about 130 m2/g tend to be more detrimentally affected by catalyst coke than catalysts having a higher surface area. Generally, the particles are usually spheroidal and have a diameter of about 1.6 to about 3.1 mm (about 1/16th- about ⅛th inch), although they may be as large as about 6.35 mm (about ¼th inch) or as small as about 1.06 mm (about 1/24th inch). In a particular reforming reactor, however, it is desirable to use catalyst particles which fall in a relatively narrow size range. A preferred catalyst particle diameter is about 1.6 mm (about 1/16th inch).
A reforming process can employ a fixed catalyst bed or a moving bed reaction vessel and a moving bed regeneration vessel. Generally, regenerated catalyst particles are fed to the reaction vessel, which typically includes several reaction zones, and the particles flow through the reaction vessel by gravity. Catalyst may be withdrawn from the bottom of the reaction vessel and transported to the regeneration vessel. In the regeneration vessel, a multi-step regeneration process is typically used to regenerate the catalyst to restore its full ability to promote reforming reactions. U.S. Pat. Nos. 3,652,231 (Greenwood et al.), 3,647,680 (Greenwood et al.) and 3,692,496 (Greenwood et al.) describe catalyst regeneration vessels that are suitable for use in a reforming process. Catalyst can flow by gravity through the various regeneration steps and then be withdrawn from the regeneration vessel and transported to the reaction vessel. Generally, arrangements are provided for adding fresh catalyst as make-up to and for withdrawing spent catalyst from the process. Movement of catalyst through the reaction and regeneration vessels is often referred to as continuous though, in practice, it is semicontinuous. By semicontinuous movement it is meant as the repeated transfer of relatively small amounts of catalyst at closely spaced points in time. For example, one batch every twenty minutes may be withdrawn from the bottom of the reaction vessel and withdrawal may take five minutes, that is, catalyst can flow for five minutes. If the catalyst inventory in a vessel is relatively large in comparison with this batch size, the catalyst bed in the vessel may be considered to be continuously moving. A moving bed system can have the advantage of maintaining production while the catalyst is removed or replaced.
Typically, the rate of catalyst movement through the catalyst beds may range from as little as about 45.5 kg (about 100 pounds) per hour to about 2727 kg (about 6000 pounds) per hour, or more.
The reaction zones of the present invention can be operated at reforming conditions, which include a range of pressures generally from atmospheric pressure about 0- about 6895 kpa(g) (about 0 psi(g)- about 1000 psi(g)), with particularly good results obtained at the relatively low pressure range of about 276- about 1379 kpa(g) (about 40- about 200 psi(g)). The overall liquid hourly space velocity (LHSV) based on the total catalyst volume in all of the reaction zones is generally about 0.1- about 10 hr−1, preferably about 1- about 5 hr−1, and more preferably about 1.5- about 4.0 hr−1.
Generally, hydrogen is supplied to provide an amount of about 1- about 20 moles of hydrogen per mole of hydrocarbon feedstock entering the reforming zone. Hydrogen is preferably supplied to provide an amount of less than about 3.5 moles of hydrogen per mole of hydrocarbon feedstock entering the reforming zone. If hydrogen is supplied, it may be supplied upstream of the combined feed exchanger, downstream of the combined feed exchanger, or both upstream and downstream of the combined feed exchanger. Alternatively, no hydrogen may be supplied to enter the reforming zone with the hydrocarbon feedstock. Even if hydrogen is not provided with the hydrocarbon feedstock to the first reaction zone, the naphthene reforming reactions that occur within the first reaction zone can yield hydrogen as a by-product. This by-product, or in-situ-produced, hydrogen leaves the first reaction zone in an admixture with the first reaction zone effluent and then can become available as hydrogen to the second reaction zone and other downstream reaction zones. This in situ hydrogen in the first reaction zone effluent usually amounts to about 0.5- about 2 moles of hydrogen per mole of hydrocarbon feedstock.
As mentioned previously, naphthene reforming reactions that are endothermic can occur in the first reaction zone, and thus the outlet temperature of the first reaction zone is often less than the inlet temperature of the first reaction zone and is generally about 316- about 454° C. (about 600°- about 850° F.). The first reaction zone contains generally about 5%- about 50%, and more usually about 10%- about 30%, of the total catalyst volume in all of the reaction zones. Consequently, the liquid hourly space velocity (LHSV) in the first reaction zone, based on the catalyst volume in the first reaction zone, is generally 0.2-200 hr−1, preferably about 2- about 100 hr−1, and more preferably about 5- about 40 hr−1. The catalyst particles are withdrawn from the first reaction zone and passed to the second reaction zone. Such particles generally have a coke content of less than about 2 wt-% based on the weight of catalyst.
The first reaction zone effluent stream can be heated in a heater, such as a gas-fired, an oil-fired, or a mixed gas-and-oil-fired heater, of a kind that is well known to persons of ordinary skill in the art of reforming. The heater may heat the first reaction zone effluent stream by radiant and/or convective heat transfer. Preferably, the first reaction zone effluent is heated in the radiant section, optimally only in the radiant section and not the convection section. The hydrocarbon stream can enter and exit the top or lower portion of the radiant section through U-shaped or inverted U-shaped tubes. Alternatively, the hydrocarbon stream can enter the top portion where the temperature is lowest in the radiant section and exit at the bottom where the temperature is hottest in the radiant section, or, conversely, enter at the bottom and exit at the top. Preferably, the hydrocarbon stream enters the top portion and exits the bottom portion of the radiant section for this and any subsequent heaters. Commercial fired heaters for reforming processes typically have individual radiant heat transfer sections for individual heaters and a common convective heat transfer section that may be heated by the flue gases from the radiant sections. Thus, this heater may be considered a second heater with the one or more convection sections being a first heater.
The first reaction zone effluent stream leaves the second heater at a temperature of generally about 482- about 560° C. (about 900- about 1040° F.). Accounting for heat losses, the heater outlet temperature is generally not more than about 5° C. (about 10° F.), and preferably not more than about 1° C. (about 2° F.), more than the inlet temperature of the second reaction zone. Accordingly, the inlet temperature of the second reaction zone is generally about 482°- about 560° C. (about 900°- about 1040° F.), preferably about 527°- about 549° C. (about 980°- about 1020° F.), and most preferably about 532° to about 543° C. (about 990°- about 1010° F.). The inlet temperature of the second reaction zone is usually at least about 33° C. (about 60° F.) greater than the inlet temperature of the first reaction zone, and may be at least about 56° C. (about 100° F.) or even at least about 83° C. (about 150° F.) higher than the first reaction zone inlet temperature. The inlet temperature of the second reaction zone is generally about 33°- about 83° C. (about 60°- about 150° F.), and preferably about 56°- about 67° C. (about 1000- about 120° F.), greater than the inlet temperature of the first reaction zone.
The desired reformate octane of the C5+fraction of the reformate is generally about 85- about 107 clear research octane number (C5+RONC), and preferably about 98- about 107 C5+RONC.
The second reaction zone generally includes about 10%- about 60%, and more usually about 15%- about 40%, of the total catalyst volume in all of the reaction zones. Consequently, the liquid hourly space velocity (LHSV) in the second reaction zone, based on the catalyst volume in the second reaction zone, is generally about 0.17- about 100 hr−1, preferably about 1.7- about 50 hr−1, and more preferably about 3.8- about 26.7 hr−1.
The second reaction effluent can pass the radiant section of a third heater, and after heating, can pass to a third reaction zone. However, one or more additional heaters and/or reactors after the second reaction zone can be omitted; that is, the second reaction zone may be the last reaction zone in the train. The third reaction zone contains generally about 25%- about 75%, and more usually about 25%- about 50%, of the total catalyst volume in all of the reaction zones. Likewise, the third reaction zone effluent can pass to the radiant section of a fourth heater and from there to a fourth reactor. The fourth reaction zone contains generally about 30%- about 80%, and more usually about 30%- about 50%, of the total catalyst volume in all of the reaction zones. The inlet temperatures of the third, fourth, and subsequent reaction zones are generally within about 1° C. (about 20° F.) of the inlet temperature of the second reaction zone.
Because the reforming reactions that occur in the second and subsequent (i.e., third and fourth) reaction zones are frequently less endothermic than those that occur in the first reaction zone, the temperature drop that occurs in the later reaction zones is often less than the drop that occurs in the first reaction zone. Thus, the outlet temperature of the last reaction zone may be about 11° C. (about 20° F.) or less below the inlet temperature of the last reaction zone, and indeed may conceivably be higher than the inlet temperature of the last reaction zone.
As previously mentioned, the last reaction zone effluent stream is cooled in the combined feed heat exchanger by transferring heat to the combined feed stream. After leaving the combined feed heat exchanger, the cooled last reactor effluent passes to a product recovery section. Suitable product recovery sections are known to persons of ordinary skill in the art of reforming. As an example, such product recovery facilities generally include gas-liquid separators for separating hydrogen and C1-C3 hydrocarbon gases from the last reactor effluent stream, and fractionation columns for separating at least a portion of the C4-C5 light hydrocarbons from the remainder of the reformate. In addition, the reformate may be separated by distillation into a light reformate fraction and a heavy reformate fraction.
In an alternative embodiment, the combined feed stream to the first reactor is heated in a radiant section of a first heater, and the third, or penultimate, reaction zone effluent can pass through one or more convection sections of at least one heater before entering a fourth, or last, reaction zone, as discussed hereinafter. The operating conditions would be similar as the embodiment discussed above. Also, it should be understood that any reaction zone in the series may have its feed heated by one or more convection sections without heating from a radiant section of a heater.
The drawings illustrate embodiments of the present invention. The drawings are presented for purposes of illustration and are not intended to limit the scope of the invention as set forth in the claims. The drawings show only the equipment and lines necessary for an understanding of the invention and do not show equipment such as pumps, compressors, heat exchangers, and valves which are not necessary for an understanding of the invention and which are well known to persons of ordinary skill in the art of hydrocarbon processing.
Referring to
The plurality of heaters 215 can include a first heater 220, a second heater 270, and a third heater 320. Generally, the first heater 220 includes a convection section 230 and a radiant section 250, the second heater 270 includes a convection section 280 and a radiant section 300, and the third heater 320 includes a convection section 330 and a radiant section 350. Each convection section 230, 280 and 330 generally includes, respectively, at least one convection tube 234, 284, and 334, and each radiant section 250, 300, and 350 generally includes, respectively, at least one burner 252 and at least one radiant tube 254, at least one burner 302 and at least one radiant tube 304, and at least one burner 352 and at least one radiant tube 354. Each convection tube 234, 284, and 334 can include, respectively, an inlet 240 and an outlet 244, an inlet 290 and an outlet 294, and an inlet 340 and an outlet 344, and each radiant tube 254, 304, and 354 can include, respectively, an inlet 260 and an outlet 264, an inlet 310 and an outlet 314, and an inlet 360 and an outlet 364. Moreover, although only one tube is discussed for each section 230, 250, 280, 300, 330, and 350 and one burner for each heater 250, 300, and 350 in the reforming unit 100, it should be understood that generally each section can include several tubes and each heater can include several burners.
During operation, a hydrocarbon stream 150 can enter the heat exchanger 200, which heats the hydrocarbon stream 150 with a hydrocarbon stream effluent 160 from the reaction zone 430. Generally, the hydrocarbon stream 150 can be referred to as a naphtha feed 154 before being subjected to the reforming reactions. Afterwards, the naphtha feed 154 may enter one or more convection sections 330, 280, and 230 to heat the feed 154 to the first reaction zone 412. Desirably, the naphtha feed 154 passes serially through the convection sections 330, 280, 230 and exits the convection section 230 directly into the first reaction zone 412 via the inlet 414. The convection sections 230, 280, and 330 generally transfer at least about 90% of the heat to the naphtha feed 154 from the at least one heater 210 before entering the first reaction zone 412. Desirably, at least about 95%, about 99% or even about 100% of the heat from the at least one heater 210 transferred to the naphtha feed 154 is from the convection sections 230, 280, and 330.
After the hydrocarbon stream 150 undergoes conversion reactions, the hydrocarbon stream 150 can exit the first reaction zone 412 via the outlet 416 into the radiant section 250 via the inlet 260. The hydrocarbon stream 150 may be heated in the radiant section 250 of the heater 220. Afterwards, the hydrocarbon stream 150 can exit via the outlet 264 to enter the second reaction zone 418.
Thus, the hydrocarbon stream 150 may be sent through the convection section 230 of the heater 220 before entering the first reaction zone 412, and the radiant section 250 of the heater 220, through which the hydrocarbon stream 150 subsequently flows, corresponds to the second reaction zone 418 in this exemplary embodiment.
Afterwards, the hydrocarbon stream 150 generally enters the second reaction zone 418 via the inlet 420 and exits via the outlet 422 into the second heater 270. Desirably, the hydrocarbon stream 150 enters the inlet 310 of the radiant section 300 to be heated for the next reaction zone 424. Afterwards, the hydrocarbon stream 150 can exit via the outlet 314 and enter the third reaction zone 424 via the inlet 426.
That being done, the hydrocarbon stream 150 can exit via the outlet 428 and enter the radiant section 350 of the third heater 320 via the inlet 360 to be heated for the fourth reaction zone 430. The hydrocarbon stream 150 can exit the heater 320 via the outlet 364 to enter the fourth reaction zone 430 via the inlet 432. Afterwards, the hydrocarbon stream 150 can exit via the outlet 434 as the hydrocarbon stream effluent 160 to heat the naphtha feed 154 in the heat exchanger 200.
Other embodiments of the present invention are depicted in
The hydrocarbon stream 150, preferably the naphtha feed 154, can be preheated in the exchanger 200 that is heated with the hydrocarbon stream effluent 160. Afterwards, the naphtha feed 154 can be heated in the convection sections 230, 280, and 330 of the plurality of heaters 215, similar as above with the convection sections 230, 280 and 330 being in reverse order with respect to the flow. Desirably, after the naphtha feed 154 exits the convection section 330 of the heater 320, the naphtha feed 154 enters the reforming reactor 450 via the inlet 454 to be converted.
Afterwards, the hydrocarbon stream 150 can exit the reforming reactor 450 via the outlet 458 and enter the radiant section 250 of the heater 220 via the inlet 260. Next, after being heated, the hydrocarbon stream 150 may exit the heater 220 via the outlet 264 and enter the second reforming reactor 460 via the inlet 464. That being done, the hydrocarbon stream 150 may undergo further conversion before exiting the reforming reactor 460 via the outlet 468.
Next, the hydrocarbon stream 150 may enter the heater 270 via the inlet 310 to be heated before exiting via the outlet 314 to undergo further reactions. After exiting the heater 270, the hydrocarbon stream 150 can enter the third reforming reactor 470 via the inlet 474 to be further reformed before exiting via the outlet 478. Subsequently, the hydrocarbon stream 150 may enter the radiant section 350 of the heater 320 via the inlet 360 for transferring heat to the hydrocarbon stream 150. That being done, the hydrocarbon stream 150 can exit via the outlet 364 to enter the fourth reforming reactor 480 via the inlet 484 for reacting the hydrocarbon stream 150 before the hydrocarbon stream 150 may exit the outlet 488. Afterwards, the hydrocarbon stream 150 can be the hydrocarbon stream effluent 160 for heating the naphtha feed 154 in the heat exchanger 200.
In another embodiment, referring to
In the embodiments described above, it is desirable to have a ratio of heater radiant sections to reaction zones or reactors of less than 1:1. Preferably, such a ratio can be 1:2, 2:3, 3:4, or 4:5 depending on the number of reaction zones. Although the flow is depicted as passing serially through the convection sections in
Moreover, although any temperature profile can be utilized with the above-described embodiments, preferably an ascending reactor inlet temperature profile is utilized. Generally, although not wanting to be bound by any theory, an ascending temperature profile lessens the variance of heat duties of a plurality of radiant sections. Such a reduced variance can improve the standardization of radiant sections in one or more heaters, thereby reducing manufacturing, installing, or refurbishing costs.
Although the embodiments described above depict heaters with their own convection section, it should be understood that reforming units described above may include one or more heaters or furnaces that have a plurality of radiant sections sharing a common convection section. Particularly, referring to
Without further elaboration, it is believed that one skilled in the art can, using the preceding description, utilize the present invention to its fullest extent. The preceding preferred specific embodiments are, therefore, to be construed as merely illustrative, and not limitative of the remainder of the disclosure in any way whatsoever.
The following prophetic examples are intended to further illustrate the subject process. These illustrations of embodiments of the invention are not meant to limit the claims of this invention to the particular details of these examples. These examples are based on engineering calculations and actual operating experience with similar processes.
Comparison Example 1 and Example 1 depict, respectively, heater duties of 4-reactor/4-heater processes and a 4-reactor/3-heater process.
In Comparison Example 1, the inlet temperatures to each reactor are equal and the fired heater with multiple radiant heater cells is used to heat up the feed and reactor effluents. The maximum process duty in the fired heater convection section available for feed heating is first estimated based on the split of duty between the convection section and radiant section and the duty available in the convection section for process heating. The following process duty or heat adsorption requirement for each heater's cell when the reactor inlet temperature is equal at 548° C. (1019° F.) is given below.
Example 1 replaces a radiant heater cell with a convection section. It is achieved by lowering the heater outlet temperature (reactor one inlet temperature in this case). In this example, only a portion of the convection section shared by the radiant heater cells is utilized to heat the feed to the first reactor (as depicted in Table 2), while the remainder is used to generate steam. The second radiant heater cell outlet temperature is also kept lower to reduce the duty requirement in the second heater cell.
The processes are moving bed processes with continuous regeneration which each reform the feedstock at the same feed rate. The LHSV, hydrogen to hydrocarbon molar ratio, reactor pressure, catalyst, C5+ RONC, catalyst distribution, and catalyst circulation rate each are the same in all the processes.
In Table 2 below, data for Comparison Example 1 are compared with data from an Example 1.
All temperatures are design temperatures which are generally 16° C. (28° F.) greater than actual predicted operating temperatures. This deviation can allow some margin of error for the predicted decrease in catalyst activity.
The example in accordance with the invention has a mean heat duty of the radiant sections of 12,400 kJ/sec. (42.36 mm btu/hr) with an unbiased standard deviation of 3,133 kJ/sec. (10.70 mm btu/hr) while Comparison Example 1 that heats the feed with four radiant cells has a mean heat duty of 11,140 kJ/sec. (38.04 mm btu/hr) and an unbiased standard deviation of 3,400 kJ/sec. (11.61 mm btu/hr). Moreover, the range of duties in Example 1 is 6,267 kJ/sec. (21.40 mm btu/hr) and the range of duties in Comparison Example 1 is 8,076 kJ/sec. (27.58 mm btu/hr).
The total fuel firing requirement for Comparison Example 1 is 77,199 kJ/sec. (263.63 mm btu/hr) at 58% radiant efficiency and 28% of the heat can be recovered by process in the convection section. An estimate of the available duty of the convection section for process heating can be made.
Available convection section duty if elimination of the first radiant heater cell:
(54.27+36.06+26.69)/0.58*(1−0.58)*0.28=23.7 mm btu/hr or 6,940 kJ/sec.
Available convection section duty if elimination of the fourth radiant heater cell:
(35.15+54.27+36.06)/0.58*(1−0.58)*0.28=25.4 mm btu/hr or 7,440 kJ/sec.
Generally, the equal reactor inlet temperature design in this comparative example is not sufficient to give enough convection section duty that replaces one of the heater cells with the same amount of catalyst loading in the reactor. However, without increasing the catalyst loading in the reactor, it is possible to eliminate the heater cell by decreasing the radiant section efficiency and increasing the bridge wall temperature or skew the reactor inlet temperature. Generally, it is not efficient to reduce the radiant section efficiency and bridge wall temperature, thus skewing the reactor inlet temperatures to change the heat adsorption requirement for each heater cell is generally more fuel and cost efficient.
Example 1 replaces the first radiant heater cell with a convection section, as discussed above. The reactor inlet temperatures are skewed to 529/542/554/554° C. (985/1008/1030/1030° F.). This helps to balance the remaining radiant section duty for different heater cells.
In Example 1, the available duty in convection section for feed heating is estimated.
(53.12+42.23+31.72)/0.58*(1−0.58)*0.28=25.8 mm btu/hr or 7,560 kJ/sec.
Thus, it is possible to reduce the radiant section efficiency slightly to below 58% or further reduce the reactor inlet temperature slightly to eliminate the first radiant heater cell. Thus, the expected fuel firing is reduced to 64,185 kJ/sec. (219.19 mm btu/hr) from 77,199 kJ/sec. (263.63 mm btu/hr) by eliminating the first radiant heater cell, which may lead to significant savings of equipment capital and the fuel.
This reduction in the range of heater duty permits the standardization of heater size, thereby reducing the cost of purchasing and/or installing such equipment in a reforming unit.
The above examples illustrate that the convection section can be used to heat the feed to a reactor, preferably the first or last reactor in a series. This can be accomplished in the above examples by properly adjusting the inlet temperature to the first reactor. Another possibility is increasing the convection section available duty for process heating to reduce the radiant section efficiency and using a higher bridge wall temperature, although this possibility is generally considered less desirable in some circumstances. Yet another possibility is adding sufficient catalyst to the reaction zones to lower the duty requirements to permit the replacement of one or more radiant sections with at least one convection section.
Without further elaboration, it is believed that one skilled in the art can, using the proceeding description, utilize the present invention to its fullest extent. Unless otherwise indicated, all parts and percentages are by weight. The entire disclosure of all cited applications, patents and publications is hereby incorporated by reference.
From the foregoing description, one skilled in the art can easily ascertain the essential characteristics of this invention and, without departing from the spirit and scope thereof, can make various changes and modifications to the invention to adapt it to various usages and conditions.
