Patent Application
20060149115
Autothermal cracking is a process for the manufacture of olefins in which a hydrocarbon feed is mixed with oxygen and passed over an autothermal cracking catalyst. The autothermal cracking catalyst is capable of supporting combustion beyond the fuel rich limit of flammability. Combustion is initiated on the catalyst surface, and the heat required to raise the reactants to the process temperature and to carry out the endothermic cracking process is generated in situ. Generally, the hydrocarbon feed and the oxygen is passed over a supported catalyst to produce the olefin product. The autothermal cracking process is described in EP 332289B; EP-529793B; EP-A-0709446 and WO 00/14035.
As with conventional furnace-based cracking, the product stream exiting the autothermal reactor is typically quenched by contact with water, and subsequently passed through a series of separation and purification steps. The product stream usually undergoes separation and purification steps to remove hydrogen, methane and CO2. The reaction products are then treated to separate methane, hydrogen and carbon monoxide before the remaining product stream is treated in order to separate a C2 containing stream from heavier hydrocarbons. The C2 containing stream is treated to separate ethylene from ethane. The remaining product stream, comprising C3 and higher hydrocarbons, may be further treated to separate propane and propylene, for example.
Ammonia absorption refrigeration (hereinafter referred to as “AMR”) systems differ from compressor-based refrigeration systems (such as a conventional C3 refrigeration system) in that they require only a relatively low-level source of enthalpy, rather than high-level energy. For example, MR systems can be driven by energy available at temperatures as low as 95° C., while compressor-based refrigeration systems typically require either superheated high-pressure steam or electricity to drive the compressors. This enthalpy is generally a waste heat source that would otherwise be lost to the atmosphere. MR systems can therefore be more energy efficient than C3 refrigeration systems. MR can be a cost-effective, energy saving process that can be used for providing moderate temperature refrigeration.
In a relatively simple MR system, an enthalpy source such as waste heat reboils an ammonia fractionator that is fed a rich ammonia aqua stream comprising a relatively high concentration of ammonia in water. The fractionator separates the strong aqua stream into a higher purity ammonia vapor overhead stream, and a bottoms stream with a lower concentration of ammonia relative to the strong aqua stream. The ammonia vapor overhead stream is condensed, typically via air or water cooling, to generate liquid ammonia refrigerant. The liquid ammonia refrigerant is then directed to the refrigerant users. As enthalpy is transferred indirectly from the material being refrigerated, the liquid ammonia refrigerant evaporates and generates ammonia refrigerant vapor. The ammonia vapor is directed to an absorber, a long with the bottoms stream from the fractionator which absorbs the ammonia vapor while releasing heat of absorption. The heat of absorption is typically removed by water cooling the absorber. Various, attempts to use MR systems to replace the propane or propylene refrigeration system in olefins manufacture have generally not been successful. There are two basic problems with using MR systems in conventional olefins plants. First, the quench water heat available in a conventional olefins plant is not available at a high enough temperature to drive the MR system. Second, conventional furnaces produce a relatively large amount of high-pressure steam by recovering high-temperature energy from the furnace flue gases. Conventional C3 refrigeration systems consume a significant fraction of this high-pressure steam, thereby helping achieve a steam balance in the olefins plant. A conventional olefins plant that utilizes an AAR system, rather than a C3 refrigeration system, would likely be significantly out of steam balance. An olefins plant based on autothermal cracking produces significantly less high-pressure steam than a conventional olefins plant, and so the use of an AAR system rather than a C3 refrigeration system within an autothermal cracking process would have less of an impact on the process steam balance. Also, in addition to providing refrigeration to cool process streams, a C3 refrigeration system in a conventional olefins plant recovers refrigeration value from cold process streams by warming them against the propylene refrigerant. Prior-art MR systems were typically not designed in such a way that they could also recover refrigeration from cold process streams. Thus, these processes have not been useful in replacing the propylene refrigeration system in a conventional olefins manufacturing process.
MR systems have been proposed for use in processes for the production of ethylene, as described in U.S. Pat. No. 4,143,521, issued to Pano et al. However, conventional cracking processes for the production of olefins, such as steam cracking furnaces, are generally operated at low pressure, and this limits the temperature of the liquid quench water that can be obtained. Typically, the maximum temperature of this water is in the range of about 80° C. to about 99° C. (approximately 180° F. to 210° F.). This relatively low-temperature water results in ammonia refrigerant being available at relatively warm temperatures for refrigeration, typically about 21° C. (70° F.), and hinders the application of MR in an olefins plant accordingly.
Use of AAR systems in ethylene plants was also suggested by D. Sohns and C. Fuge, “Reducing Consumption of Energy Is Possible in Olefin Plants,” Oil & Gas Journal, Sep. 13, 1976, pp 72-77. In this reference, the heat source to drive the MR system was quench oil, a stream that is prone to fouling and is not available in all olefins plants. The authors state that the quench water stream in a conventional olefins plant is of little utility for providing refrigeration.
Although the temperature of liquid water obtainable can, in theory, be increased by operating the cracking process at higher pressures, this is not desirable for conventional furnace-based cracking processes because the optimum pressure for the furnace-based cracking reaction is generally less than about 2 bar. Thus, conventional furnace-based cracking processes typically use a C3 refrigeration system, or a variation thereof to provide refrigeration at temperatures between ambient and about −45° C. C3 refrigeration systems, as described in U.S. Pat. No. 6,637,237 are generally powered by high pressure steam generated in the furnace-based cracking process. Although they are less energy efficient than AAR systems, this has not been a significant concern because conventional furnace-based cracking processes produce a large amount of high pressure steam which can be used for the C3 refrigeration systems.
It would be highly desirable to make the ammonia refrigeration available at significantly lower temperatures, for example down to about −45° C. (−50° F.), while also recovering refrigeration value from the relatively low temperature process streams available in processes that produce olefins.
Surprisingly, we have now found that olefins may be advantageously produced by autothermal cracking of hydrocarbons at relatively high pressures, where the water from quenching of the autothermal cracking product stream is utilized to drive an MR refrigeration system for purification of the product stream to produce said olefins. In particular, a beneficial feature of the present invention is that the ammonia refrigeration is made available at a lower temperature than the prior art processes, and that the MR system can completely replace the conventional C3 refrigeration system within olefins manufacturing plants based on the relatively high-pressure autothermal cracking of hydrocarbons.
One aspect of this invention is an autothermal cracking process for production and recovery of olefins. The process comprises feeding a substantially hydrocarbon feedstock and an oxygen-containing gas to an autothermal cracker to provide a hydrocarbon product stream comprising olefins. A waste enthalpy source generated by said autothermal cracking process is used to at least partially drive an ammonia absorption refrigeration system to provide chilling for at least one process stream in the separation and/or purification of olefins from the hydrocarbon product stream.
Another aspect of this invention is an ammonia absorption refrigeration process comprising at least one enthalpy source selected from the group consisting of: quench water generated through the cooling of cracked gases from an autothermal cracking reactor; steam generated through the cooling of cracked gases from an autothermal cracking reactor; sufficiently warm streams derived from processes which utilize the ethylene produced from the autothermal cracking process; and sufficiently warm streams from other chemical or refinery process units located near an autothermal cracking reactor.
This invention describes using an AAR system to recover olefins, including ethylene, from a cracked gas stream which is produced by an autothermal cracking reactor. There are many design options for the recovery of ethylene from a cracked gas stream. The process of the present invention may be used to convert both liquid and gaseous hydrocarbons into olefins. Suitable liquid hydrocarbons include naphtha, gas oils, vacuum gas oils and mixtures thereof. Preferably, however, gaseous hydrocarbons such as ethane, propane, butane and mixtures thereof are employed.
A substantially hydrocarbon feedstock to the process is shown as stream 1. As used herein, the term “substantially hydrocarbon feedstock” refers to a hydrocarbon feedstock that generally comprises hydrocarbons consisting essentially of ethane, ethylene, propane, propylene, butane, butylene, diene and acetylenic compounds, and hydrocarbon impurities. It is combined with an oxygen-containing gas, shown as stream 2. Suitably, the oxygen-containing gas is molecular oxygen, air and/or mixtures thereof. The oxygen-containing gas may be mixed with an inert gas such as nitrogen or argon. Optionally, a recycle stream, shown as stream 3 and a hydrogen-containing stream, shown as stream 4 may enter the autothermal reactor 5. In the autothermal reactor 5, streams 1 through 4 can be preheated and are reacted to form a hot cracked gas, shown as stream 6. Preferably, the substantially hydrocarbon feedstock and oxygen-containing gas are fed to the autothermal reactor 5 at a ratio of hydrocarbon to oxygen-containing gas of about 5 to about 16 times, preferably about 5 to about 13.5 times, more preferably about 6 to about 10 times, the stoichiometric ratio of hydrocarbon to oxygen-containing gas required for complete combustion of the hydrocarbon to carbon dioxide and water.
The autothermal cracking catalyst is capable of supporting combustion beyond the fuel rich limit of flammability. The catalyst usually comprises a Group VIII metal as its catalytic component. Suitable Group VIII metals include platinum, palladium, ruthenium, rhodium, osmium and iridium. Rhodium, and more particularly, platinum and palladium are preferred. Typical Group VIII metal loadings range from about 0.01 to about 100 weight percent, preferably, between about 0.01 to about 20 weight percent, and more preferably, from about 0.01 to about 10 weight percent based on the total dry weight of the catalyst.
Where a Group VIII catalyst is utilized, it is preferably utilized in combination with a catalyst promoter. The promoter may be a Group IIIA, IVA, and/or VA metal. Alternatively, the promoter may be a transition metal; the transition metal promoter being a different metal to that which may be employed as the Group VII transition metal catalytic component.
Preferred Group IIIA metals include Al, Ga, In and TI. Of these, Ga and In are preferred. Preferred Group IVA metals include Ge, Sn and Pb. Of these, Ge and Sn are preferred. The preferred Group VA metal is Sb. The atomic ratio of Group VII B metal to the Group IIIA, IVA or VA metal may be about 1: about 0.1-50.0, preferably, about 1: about 0.1-12.0.
Suitable metals in the transition metal series include those metals in Group IB to VII of the Periodic Table. In particular, transition metals selected from Groups IB, IIB, VIIB, VIIB and VII of the Periodic Table are preferred. Examples of such metals include Cr, Mo, W, Fe, Ru, Os, Co, Rh, Ir, Ni, Pt, Cu, Ag, Au, Zn, Cd and Hg. Preferred transition metal promoters are Mo, Rh, Ru, Ir, Pt, Cu and Zn. The atomic ratio of Group VIII metal to transition metal promoter may be about 1: about 0.1-about 50.0, preferably, about 1: about 0.1-about 12.0.
Preferably, the catalyst comprises only one promoter selected from Group IIIA, Group IVA, Group VB and the transition metal series. For example, the catalyst may comprise a metal selected from rhodium, platinum and palladium and a promoter selected from the group consisting of Ga, In, Sn, Ge, Ag, Au or Cu. Preferred examples of such catalysts include Pt/Ga, Pt/In, Pt/Sn, Pt/Ge, Pt/Cu, Pd/Sn, Pd/Ge, Pd/Cu and Rh/Sn. The Rh, Pt or Pd may comprise between about 0.01 and about 5.0 weight percent, preferably, between about 0.01 and about 2.0 weight percent, and more preferably, between about 0.05 and about 1.0 weight percent of the total weight of the catalyst. The atomic ratio of Rh, Pt or Pd to the Group IIIA, IVA or transition metal promoter may be about 1: about 0.1-about 50.0, preferably, about 1: about 0.1-about 12.0. For example, atomic ratios of Rh, Pt or Pd to Sn may be about 1:0.1 to about 50, preferably, about 1:0.1-about 12.0, more preferably, about 1: about 0.2-about 3.0 and most preferably, about 1: about 0.5-about 1.5. Atomic ratios of Pt or Pd to Ge, on the other hand, may be about 1: about 0.1 to about 50, preferably, about 1: about 0.1-about 12.0, and more preferably, about 1: about 0.5-about 8.0. Atomic ratios of Pt or Pd to Cu may be about 1: about 0.1-about 3.0, preferably, about 1: about 0.2-about 2.0, and more preferably, about 1: about 0.5-about 1.5.
Alternatively, the promoter may comprise at least two metals selected from Group IIIA, Group IVA and the transition metal series. For example, where the catalyst comprises platinum, the platinum may be promoted with two metals from the transition metal series, for example, palladium and copper. Such Pt/Pd/Cu catalysts may comprise palladium in an amount of about 0.01 to about 5 weight percent, preferably, about 0.01 to about 2 weight percent, and more preferably, about 0.01 to about 1 weight percent based on the total weight of the dry catalyst. The atomic ratio of Pt to Pd may be about 1: about 0.1-about 10.0, preferably, about 1: about 0.5-about 8.0, and more preferably, about 1: about 1.0-about 5.0. The atomic ratio of platinum to copper is preferably about 1: about 0.1-about 3.0, more preferably, about 1: about 0.2-about 2.0, and most preferably, about 1: about 0.5-about 1.5. Where the catalyst comprises platinum, it may alternatively be promoted with one transition metal, and another metal selected from Group IIIA or Group IVA of the periodic table. In such catalysts, palladium may be present in an amount of about 0.01 to about 5 weight percent, preferably, about 0.01 to about 2.0 weight percent, and more preferably, about 0.05-about 1.0 weight percent based on the total weight of the catalyst. The atomic ratio of Pt to Pd may be about 1: about 0.1-about 10.0, preferably, about 1: about 0.5-about 8.0, and more preferably, about 1: about 1.0-about 5.0. The atomic ratio of Pt to the Group IIIA or IVA metal may be about 1: about 0.1-about 60, preferably, about 1: about 0.1-about 50.0. Preferably, the Group IIIA or IVA metal is Sn or Ge, most preferably, Sn.
For the avoidance of doubt, the Group VIII metal and promoter in the catalyst may be present in any form, such as a metal, or in the form of a metal compound, such as an oxide.
The catalyst may be unsupported, such as in the form of a metal gauze, but is preferably supported. Any suitable support may be used such as ceramic or metal supports, but ceramic supports are generally preferred. Where ceramic supports are used, the composition of the ceramic support may be any oxide or combination of oxides that is stable at high temperatures of, for example, between about 600° C. and about 1200° C. The support material preferably has a low thermal expansion co-efficient, and is resistant to phase separation at high temperatures.
Suitable ceramic supports include corderite, lithium aluminum silicate (LAS), alumina (α-Al2O3), yttria stabilized zirconia, alumina titanate, niascon, and calcium zirconyl phosphate. The ceramic supports may be wash-coated, for example, with γ-Al2O3.
The catalyst capable of supporting combustion beyond the fuel rich limit of flammability may be prepared by any method known in the art. For example, gel methods and wet-impregnation techniques may be employed. Typically, the support is impregnated with one or more solutions comprising the metals, dried and then calcined in air. The support may be impregnated in one or more steps. Preferably, multiple impregnation steps are employed. The support is preferably dried and calcined between each impregnation, and then subjected to a final calcination, preferably, in air. The calcined support may then be reduced, for example, by heat treatment in a hydrogen atmosphere.
The hydrocarbon-containing feedstock is passed over the autothermal cracking catalyst at a gas hourly space velocity of greater than about 10,000 h−1, preferably above about 20,000 h−1 and most preferably, greater than about 100,000 h−1. It will be understood, however, that the optimum gas hourly space velocity will depend upon the pressure and nature of the feed composition.
Additional feed components may be co-fed into the autothermal cracking reactor 5, such as hydrogen, carbon monoxide, carbon dioxide or steam. Preferably, the reactant mixture of hydrogen co-fed with the hydrocarbon-containing feedstock and oxygen-containing gas into the autothermal cracking reactor 5, and preheated prior to contact with the catalyst. Suitably, the molar ratio of hydrogen to oxygen-containing gas is in the range about 0.2 to about 4. Hydrogen co-feeds are advantageous because, in the presence of the catalyst, the hydrogen combusts preferentially relative to the hydrocarbon, thereby increasing the olefin selectivity of the overall process. Generally, the reactant mixture is preheated to temperatures below the auto ignition temperature of the reactant mixture.
A heat exchanger may be employed to preheat the reactant mixture prior to contact with the catalyst. The use of a heat exchanger may allow the reactant mixture to be heated to high preheat temperatures such as temperatures at or above the autoignition temperature of the reactant mixture. The use of high pre-heat temperatures is beneficial in that less oxygen reactant is required which leads to economic savings. Additionally, the use of high preheat temperatures can result in improved selectivity to olefin product. It has also been found that the use of high preheat temperatures enhances the stability of the reaction within the catalyst thereby leading to higher sustainable superficial feed velocities, and also reduces the thermal gradient experienced across the catalyst.
The autothermal cracking process may suitably be carried out at a catalyst exit temperature in the range of about 600° C. to about 1200° C., preferably, in the range of about 850° C. to about 1050° C. and, most preferably, in the range of about 900° C. to about 1000° C. To avoid further reactions taking place, the product stream is preferably is cooled to between about 600° C. and about 750° C. within 20 milliseconds of formation to form the hot cracked gas stream 6. Advantageously, if the autothermal cracking process is operated at a pressure of greater than about 20 barg, the products are cooled to between about 600° C. and about 750° C. within 10 milliseconds of formation. The autothermal cracking of the present invention is operated at a pressure of greater than about 5 barg. Preferably the autothermal cracking process is operated at a pressure of between about 5 to about 40 barg and preferably between about 10 to about 30 barg.
The hot cracked gas, shown as stream 6 typically contains ethylene, methane, hydrogen, carbon monoxide, carbon dioxide, ethane, and hydrocarbons heavier than ethane. The hot cracked gas stream 6 is cooled to by indirect heat exchange with boiler feed water in the primary cooling system, shown as step 7. This system generally has one or more heat exchangers and produces high-pressure steam and a cooled cracked gas, shown as stream 8.
The cooled cracked gas stream 8 is further cooled in a quench section, shown as step 9. The cooling processes employed, which are well known to those skilled in the art of ethylene manufacture, typically involve at least a contacting vessel in which the direct contact of the cracked gas with a circulating cooled water stream takes place. An additional step of direct contact with a circulating cooled hydrocarbon stream may optionally be employed within the quench section 9. The contacting operation generates a cooled cracked gas, shown as stream 10 and a warmed quench water stream. This warmed quench water is typically cooled in one or more exchangers and re-circulated to the direct contact vessel (the quench water circuit is not shown in
Stream 10 can be directed into a compression and contaminant removal step 11. Within this step the cooled cracked gas stream 10 is compressed to a pressure suitable for the downstream separation section, and contaminants are removed. For example, an amine or caustic scrubber may be employed to remove carbon dioxide and other acid gases from the cracked gas. Water is also typically removed from the cracked gas by condensation and/or adsorbent driers. If the autothermal cracking reaction is carried out at a sufficiently high pressure, the cracked gas may not need compression within step 11. In this case only contaminant removal would take place within step 11.
The high-pressure, essentially contaminant free cracked gas, shown as stream 12 is then chilled and partially condensed in exchanger 13. In practice, exchanger 13 would typically represent a series of exchangers and vapor/liquid separation vessels in which the cracked gas is progressively cooled and partially condensed by various cold process streams and progressively colder levels of external refrigeration. In a typical olefins plant this refrigeration would be supplied by a propylene refrigeration system and an ethylene or mixed refrigerant refrigeration system. The chilled cracked gas, shown as stream 14, typically enters the demethanizer column 15 at a temperature less than about −35° C. For simplicity stream 14 is depicted in
The demethanizer column 15 separates the methane and lighter components from the ethylene and heavier components in the cracked gas. The column is refluxed using partial condenser 16. The net overhead product of the demethanizer, stream 17, contains primarily methane, hydrogen, and carbon monoxide and little, if any, ethylene and heavier components. Stream 17 can be directed to hydrogen and/or CO recovery sections if desired, or used as fuel. The demethanizer 15 is reboiled with exchanger 18. The bottoms product of the demethanizer, stream 19, contains primarily ethylene and heavier components and little, if any, methane and lighter components.
Stream 19 enters the deethanizer column 20, which separates the ethane and lighter components from those heavier than ethane. The deethanizer 20 is refluxed with condenser 21 and reboiled with exchanger 22. The bottoms product, shown as stream 23, contains primarily hydrocarbons heavier than ethane. Stream 23 can be further treated to recover one or more of the heavy hydrocarbons if desired. The deethanizer net overhead product, shown as stream 24, enters an acetylene conversion system, shown as step 25. This system typically contains a number of exchangers and reactors arranged such that the deethanizer overhead stream 24 is first heated, then essentially all of the acetylene in the stream is reacted with an external hydrogen stream 26, whereupon the stream is cooled again.
The essentially acetylene free, cooled stream 27 enters a C2 splitter column 28, which purifies the ethylene sufficiently to be sold as a commercial product. The C2 splitter is refluxed with condenser 31 and reboiled with exchanger 29. The net overhead product is typically reheated in exchanger 32 and then withdrawn as the final purified ethylene product, shown as stream 33. As is well-known by those skilled in the art, a pasteurizing section can be utilized on the top section of the C2 splitter column 28. In this case the final ethylene product would be withdrawn as a liquid from an intermediate stage of column 28, and the top vapor stream would serve as a vent for light gases. The bottoms product 30 contains primarily ethane and can be recycled to the reactor section or sold as a product.
It has been discovered that the autothermal cracking based olefins production process represented in
This synergy is exhibited when the autothermal cracking reactor 5 and therefore the quench step 9 are operated at a relatively high pressure, above about 5 bar absolute. In this case the quench water from step 9 can be recovered at a sufficiently high temperature to efficiently contribute to the operation of an AAR system.
A second portion of stream 100 enters the ammonia generator column 107 at a relatively high location as stream 108. Optionally, a third portion of stream 100, stream 109, can be heated in exchanger 110 against the partially cooled quench water stream 111. The resulting heated stream 112 enters the ammonia generator column 107 at a middle location. At least a part of the cooled quench water stream 113 can be returned directly to step 9 or it can be further cooled before returning to step 9.
Many other types of arrangements for heating the rich ammonia aqua stream 100 and feeding it to the ammonia generator column can be envisioned by those skilled in the art of process design and optimization. The optimal column feed design will depend on the available utilities as well as economic (e.g. capital cost) and operational factors. The invention includes all such design variations, including, but not limited to, multiple feed locations, multiple ammonia concentrations, and multiple levels of preheat.
The ammonia generator column 107 is reboiled with medium-pressure steam (typically 6-25 bar) in exchanger 114. Other high-temperature, preferably waste enthalpy sources, could be used to reboil column 107. The net bottom product stream 115 consists of relatively lean ammonia aqua (enriched in water relative to the rich ammonia aqua stream 100). It will typically consist of water with less than 50 weight percent ammonia, and preferably less than 20 weight percent ammonia. It is cooled in exchanger 102 by contact with the rich ammonia aqua stream 101, optionally further cooled against cooling water in exchanger 116, and then directed to the absorption section as stream 117.
The gross overhead product of the ammonia generator, stream 118, consists of essentially pure ammonia and is condensed in exchanger 119. This stream will typically consist of at least 95 weight percent ammonia, preferably at least 99 weight percent ammonia. At least a portion of this condensed ammonia is commonly vaporized by indirect transfer of heat from one or more streams within the autothermal cracking olefin production process to said condensed ammonia. Indirect heat transfer means that the refrigerant is not in direct contact with the material being cooled, but rather, the refrigerant and the material being cooled are on opposite sides of a heat transfer surface. The liquid ammonia product is sent to an ammonia accumulator 120. A portion of the liquid ammonia is returned to the top of the ammonia generator as reflux liquid. The remainder, the net liquid overhead product of the ammonia generator, is withdrawn at around 15 barg as stream 121. A portion of stream 121, stream 122 can be subcooled in exchanger 123 against a cold stream from the olefins separation process, for example a stream at a temperature below about 10° C. For example and with reference to
The pressure of stream 128 is reduced to a lower pressure, for example 0.5 to 1.0 barg through valve 129 or some other pressure-reducing means. The flashed and at least partially liquid stream 130 is directed to exchanger 131 where it is vaporized to provide refrigeration to as low as about −45° C. to the olefins separation and purification process. Exchanger 131 will generally represent a number of individual exchangers to provide refrigeration to discrete points in the olefins separation and purification process. For example and with reference to
The vaporized low-pressure ammonia stream 132 is reheated in exchanger 125. The resulting heated stream 133 is then split into two portions, stream 134 and 138. Stream 134 is directed to the ambient temperature absorber 135. In this absorber 135 the ammonia is absorbed into the lean ammonia aqua stream 117. The absorption of ammonia into water is exothermic. Thus, cooling water is used to cool the absorber to drive the absorption. The heat of absorption in absorber 135 is removed with cooling water or some other ambient-temperature cooling medium. The intermediate-concentration aqua stream 136 from absorber 135 is directed to the sub-ambient absorber 137 where it is contacted with the remainder of the vaporized ammonia, stream 138. The heat of absorption in absorber 137 is removed by a sub-ambient cooling medium. For example and with reference to
Many other types of arrangements for absorbing the heated ammonia stream 133 into the lean ammonia aqua stream 117 can be envisioned by those skilled in the art. The optimal absorber system design will depend on the available ambient and sub-ambient cooling sources as well as economic (e.g. capital cost) and operational factors. The invention includes all such design variations, including, but not limited to, multiple absorption steps, absorption at multiple pressures, and multiple series/parallel arrangements of individual absorption steps
The embodiment of
In the process of
The resulting rich ammonia aqua stream 211 is pumped and combined with the rich ammonia aqua stream 100 and subsequently fed to the ammonia generator column 107 as described above. Alternatively, the rich ammonia aqua stream 211 could be fed directly in one or more portions to the ammonia generator column 107. In general, depending on the design chosen, stream 211 could also be fed to the lower-pressure absorbers 135 and/or 137, combined with the rich aqua from absorbers 135 and/or 137, or fed separately to the ammonia generator column 107. All such variations are encompassed within the concept of this invention.
One aspect of this invention and its integration with an autothermal cracking process to produce olefins is that one or more sub-ambient temperature streams from the olefins recovery and purification process are heated in and thereby provide cooling duty to one or more parts of the ammonia absorption process of this invention. In each case the heating of these olefins related process streams improves the performance or efficiency of the AAR system of this invention. As described in
As further described in
In the present invention, streams within the olefins recovery and purification process are condensed and optionally desuperheated by exchange with ammonia refrigerant. The ammonia evaporation temperature is typically in the range of about 10° C. to about −45° C. Ammonia is vaporized by heat indirectly transferred from the relevant steps discussed in
In a typical olefins recovery and purification process refrigeration is required at temperatures below that which can be provided by a C3 or MR refrigeration system. In practice, refrigeration at these colder temperatures is typically provided by a separate ethylene refrigeration system or a mixed refrigeration system. Ammonia evaporation temperatures as low as about −45° C. can readily be achieved in an MR system comprising an ammonia refrigerant. This is sufficient to condense ethylene refrigerant at typical ethylene refrigeration compressor discharge pressures. It is also sufficient to provide condensing duty to a mixed refrigeration stream. Thus, the combination of MR and an ethylene refrigeration system, or MR and a mixed refrigeration system is sufficient to provide all of the net refrigeration needs within the ethylene recovery and purification process.
Advanced MR cycles, including multi-stage absorption refrigeration systems, multiple-lift refrigeration cycles, advanced absorption vapor exchange GAX cycles, and multiple effect absorption cycles, as described in U.S. Pat. No. 5,097,676, U.S. Pat. No. 5,966,948, Erickson and Tang, “Evaluation of Double-Lift Cycles for Waste Heat Powered Refrigeration,” Intl. Absorption Conf., Montreal, Canada, Sept. 17-22 (1996), Erickson, Potnis, and Tang, “Triple Effect Absorption Cycles,” Proc. Intersoc. Energy Convers. Eng. Conf. (1996), 31st, 1072-1077, Rane and Erickson, “Advanced absorption cycle: vapor exchange GAX,” Am. Soc. Mech. Eng. (1994) 25-32, and Richter, “Multi-Stage Absorption Refrigeration Systems,” Journal of Refrigeration, September/October 1962, are hereby incorporated by reference. Advanced MR cycles can use less heat and lower temperature heat sources while providing refrigeration at lower temperatures than simpler MR processes. Furthermore, the advanced MR cycles can accommodate refrigeration at multiple temperature levels and heat sources at multiple levels. Advanced MR cycles can have multiple absorbers and multiple ammonia fractionators.
It is generally most preferred to completely forego or replace the propane or propylene refrigeration circuit with an MR circuit, since this allows complete elimination of the energy intensive C3 compressor, condenser, flash drums, and other equipment associated with the circuit, as well as elimination of the utilities consumption associated with running the C3 compressor. The evaporators are generally the interface between the refrigeration circuit and the process. The process stream being cooled is on the hot side of the evaporator, and evaporating refrigerant is on the cold side of the evaporator. Thus, when using an MR system to replace a C3 refrigeration circuit, the evaporators retain their function and boiling ammonia refrigerant replaces boiling C3 refrigerant on cold side of the evaporator.
Replacing the C3 refrigeration circuit of an autothermal cracking process with an MR system will generally result in lower energy consumption and higher waste heat utilization. The major power input to conventional C3 refrigeration cycles is in the form of electricity or high-pressure steam used to power the compressor motor. The major power to an MR unit is the waste enthalpy source used to preheat feeds to the ammonia fractionator and/or to reboil the ammonia fractionator. The waste enthalpy source is essentially free energy, since it is otherwise lost to the environment via air or water cooling. Thus, replacing the C3 refrigeration cycle with an MR refrigeration cycle generally leads to savings of at least the electricity or steam required to power the drivers of the propane or propylene compressors, since only a small amount of electricity or steam is required to power the drivers of the pumps associated with the MR.
The use of an MR system allows the elimination of the conventional C3 system, and provides a more energy efficient refrigeration system. Thus, the MR system may be used to provide all net refrigeration duty between about −45° C. and ambient temperature for the autothermal cracking olefins plant. In addition, the use of an MR system utilizes the waste enthalpy source (from the quench water) that would otherwise be lost and the use of an AAR system reduces the high pressure steam requirement for the overall olefins plant. This is particularly beneficial for autothermal cracking processes since such processes produce significantly less high pressure steam than conventional furnace-based crackers.
For this invention, it is preferable that at least portion of the waste enthalpy source used in the MR fractionator is derived from a heat source available from the autothermal cracking process, from a unit that produces feed for the autothermal cracking process, or from a unit that is located near the autothermal cracking process. Suitable sources of heat to the MR ammonia fractionator are those that are available at a supply temperature of at least 95° C., and preferably at least 110° C. for best results. Higher waste enthalpy source stream temperatures are preferred since they generally lead to higher MR process efficiency.
One suitable waste enthalpy source on the autothermal cracking unit is the quench water generated through the cooling of cracked gases from an autothermal cracking reactor.
Another suitable waste enthalpy source could be saturated high-pressure steam generated through the cooling of cracked gases from an autothermal cracking reactor, or waste low- or medium-pressure steam from the autothermal cracking process.
Still another suitable waste enthalpy source may be derived from processes which utilize the ethylene produced from the autothermal cracking process, such as polyethylene or ethylene oxide manufacture.
Waste enthalpy sources for the AAR are not limited to those described. They can also include heat sources available on other nearby chemical or refinery process units, and steam which may be available from these units or site utilities units. The use of a waste heat enthalpy source from autothermal cracking process streams, heat from processes that produce a feed stream for the autothermal cracking process, or heat sources available on other chemical or refinery process units located near the autothermal cracking process provides synergy between the autothermal cracking process and these other processes.
The process of the present invention results in substantial benefits over alternative autothermal cracking processes. One benefit is that utilizing AAR for autothermal cracking processes in accordance with the present invention allows for the elimination of a propane or propylene refrigeration loop commonly used in conventional autothermal cracking processes. This eliminates the expensive C3 compressor, condenser, flash drums, and other equipment associated with the circuit, as well as elimination of the capitalized utilities associated with running the C3 compressor. The cost of suitable AAR in accordance with the present invention is substantially lower than conventional propane or propylene vapor recompression systems.
Another benefit of utilizing MR within an autothermal cracking process is that replacing the C3 refrigeration cycle will lead to energy savings approximately equal to the electricity or steam required to power the drivers of the propane or propylene compressors of a conventional autothermal cracking process, since a relatively smaller amount of electricity and/or steam is required to power the drivers of the pumps associated with MR system and to provide other process heat required by the MR system.
Another benefit is that utilizing MR within an autothermal cracking process in accordance with the present invention consumes waste heat enthalpy from the autothermal cracking process for preheating the feed to the ammonia fractionator. Waste enthalpy sources are essentially a free enthalpy source, since it is otherwise lost to the environment via air or water cooling.
Another benefit is that the MR system is driven by pumps for conveying liquids as compared to refrigeration compressors for conveying gas. Refrigeration compressors are far more costly and require more energy to operate than pumps that convey liquid. Since compression often results in an elevation in the temperature of the compressed gas due to compressor inefficiency, inevitably, additionally cooling utilities are required and energy lost.
Another benefit is that utilizing MR for autothermal cracking processes in accordance with the present invention also reduces greenhouse gas emissions. The use of waste heat powered MR in autothermal cracking processes generally results in a substantial reduction in electricity or high-pressure steam consumption from the overall replacement of vapor recompression refrigeration compressors with MR pumps. Reducing electricity or high-pressure steam consumption generally leads to lower CO2 emissions, since incremental electricity or high-pressure steam most often derives from fossil fuel fired power plants or plant furnaces.
The process has been described for the purposes of illustration only in connection with certain embodiments. However, it is recognized that various changes, additions, improvements, and modifications to the illustrated embodiments may be made by those persons skilled in the art, all falling within the scope and spirit of the invention.
This example describes the process of the present invention for recovering olefins, and in particular ethylene, from a mixed hydrocarbon stream derived from the effluent of an autothermal cracking reactor. The ammonia absorption process of this example was simulated using a commercially available process simulation package. The process simulated in the example is identical to the embodiment of
Exchanger 131, the net refrigeration duty supplied by the ammonia absorption refrigeration system of this invention, is depicted in
If the high-temperature quench water stream 104 were not used to pre-heat the rich ammonia-water solution in exchanger 103, an additional 12 MW of steam thermal energy would be required in the ammonia generator, either as additional medium-pressure steam in reboiler 114, or as low-pressure (5 bar) steam in exchanger 105. This corresponds to about 20,500 kg/hr of medium- or low-pressure stream. Therefore the use of waste heat in the quench water allows for a significant savings in higher-value steam energy.
This example describes the process of the present invention for recovering olefins, and in particular ethylene, from a mixed hydrocarbon stream derived from the effluent of an autothermal cracking reactor king reactor. In this example both low-temperature and intermediate-temperature ammonia refrigeration circuits are used, and there is direct recuperation of ammonia refrigerant in exchanger 213. The process simulated in this example is identical to the preferred embodiment of
As in Example 1, exchanger 131 is depicted in
This example demonstrates the flexibility of the process of this invention to providing refrigeration at a number of temperatures, and the ability to recuperate refrigeration using cold process streams, for example in exchangers 123 and 213 and absorber 137.
In this example, if the high-temperature quench water stream 104 were not used to pre-heat the rich ammonia-water solution in exchanger 103, an additional 15-16 MW of steam thermal energy would be required in the ammonia generator, either as additional medium-pressure steam in reboiler 114, or as low-pressure (5 bar) steam in exchanger 105. This corresponds to about 26,000-27,000 kg/hr of medium- or low-pressure steam. Therefore the use of waste heat in the quench water allows for a significant savings in higher-value steam energy.
