AU2012227278B2 - Multiple reactor chemical production system - Google Patents

Abstract

The present invention is a multiple reaction set for the production of chemicals by equilibrium limited reactions utilizing plate-type or extended surface heat exchangers. The heat exchangers effectively cool the reaction products in order to condense the methanol contained within the reaction products for separation, and also to warm incoming feed reactants prior to entrance of the reactants into a reactor utilized for the production of methanol. The various reactors, heat exchangers, and separators can be formed as separated zones within the enclosed vessels, thereby eliminating the need for separately constructed reactors, heat exchangers, and separators. Multi-stream plate-type of extended surface heat exchangers can be utilized to allow efficient cooling and methanol separation. The multiple reaction set can also be used for the recovery of methanol from a waste or purge gas stream utilizing multiple reactors, multiple plate-type or extended surface heat exchangers and multiple separators as a substitute for or in conjunction with a conventional methanol synthesis loop. WO 20071096699 PCT/IB2006/003279 Rxn Rxn Rx Rxn. --- d/Effluen hetpur changers 7 Spn Sepn Divided wall reactor vessel Integrated feed-effluent heat exchanger and divided wall gas liquid separation vessel OFeed- ifia nt xch n sctio Sepn Sepa Sepn Kn ok out secti n Methanol

Description

AUSTRALIA Regulation 3.2 Patents Act 1990 Complete Specification Standard Patent APPLICANT: Heatric Invention Title: MULTIPLE REACTOR CHEMICAL PRODUCTION SYSTEM The following statement is a full description of this invention, including the best method of performing it known to me: 1 MULTIPLE REACTOR CHEMICAL PRODUCTION SYSTEM 2 3 CROSS-REFERENCE TO RELATED APPLICATIONS 4 This is a divisional application of U.S. patent application Ser. No. 5 12/067,608, filed on Aug. 6, 2008 now U.S. Pat. No. 8,075,856, which claimed 6 priority under 35 U.S.C. §119(e) to U.S. Provisional Patent Application Serial No. 7 60/720,330, filed on September 23, 2005 the entirety of which is expressly 8 incorporated by reference herein. 9 10 BACKGROUND OF THE INVENTION 11 It is rare in a chemical process that the process uses stoichiometric 12 amounts of reactants with essentially complete conversion in a simple reactor. 13 Hence, where the reagents constitute a significant cost of the process, the 14 unreacted material is often recycled to the reactor, usually after some physical 15 separation of the desired product from the un-reacted material. Sometimes this 16 separation can be achieved internally within the reactor, for example, where the 17 reactants are gaseous, the product is in liquid form at reaction conditions and is 18 withdrawn continuously, and a stirred tank reactor with gas induction impellers is 19 utilized. In this particular situation, the physical separation and the recycle occurs 20 within the reactor vessel. 21 22 Alternatively, the separation and recycle of reactant can take place external 23 to the reactor. One example of this configuration would be a plug flow gas phase 24 reactor where the product can be condensed from the gas phase by cooling. The 25 unreacted gases may then be re-compressed, and at least partially returned to the 26 inlet of the reactor, perhaps after other conditioning, such as purification or 27 chemical separation. 28 There are several reasons why the amounts of reactants utilized in the 29 reactor to form the desired end product are rarely stoichiometric. It could be that 30 vapor pressure limitations require a non-stoichiometric reaction. For example, in 31 high-pressure, gas-phase hydrogenations of a high boiling organic the hydrogen 32 will be present in large excess even though a high single pass conversion of the 33 WO 2007/096699 PCT/IB2006/003279 1 reactant is theoretically possible. An alternate reason would be that the reaction is 2 equilibrium limited. For example in acid catalysed esterifications the alcohol is 3 frequently in excess to achieve high conversion of the acid. 4 5 While it may be possible to achieve high conversion of the reactants in an 6 equilibrium limited reaction to the desired end product economically utilizing a 7 large excess of reactants, an alternate possibility is the removal of one of the 8 reaction products. For example, gas stripping could remove the water from an 9 esterification reaction to continually move the conditions in the reactor out of or 10 away from equilibrium and thereby drive the reaction toward full conversion. 11 12 However, where a product cannot be removed in-situ to drive the 13 equilibrium toward full conversion, then a high overall conversion is likely to be 14 achieved only by separation of the product from the reactant mixture and 15 subsequent recycle of the un-reacted material to the reactor. 16 17 Additionally, even where an equilibrium reaction has certain conditions or 18 aspects that are favorable for high conversion to the desired end product, the 19 kinetics of the reaction may suggest that a higher overall production rate or better 20 process economics can be achieved by running a reactor at conditions favoring a 21 relatively low conversion of the reactants and then recycling the un-reacted 22 material after physical separation of the product. Some exemplary reactions or 23 processes of this type where conversion of the reactants to the desired product is 24 only partial, and in which significant quantities of un-reacted material remain that 25 can be recycled to the reactor after physical separation of the product include the 26 reaction of synthesis gas to methanol, di-methyl ether, mixtures thereof, Fischer 27 Tropsch waxes, and ammonia. 28 29 Using the equilibrium limited reaction for the production of methanol as 30 an example, because methanol is one of the largest by volume chemicals produced 31 in the world today, the conversion to methanol is typically carried out in a two 2 WO 20071096699 PCT/IB2006/003279 I step process. In a first step, methane is reformed with water or partially oxidized 2 with oxygen to produce carbon monoxide and hydrogen, with some carbon 3 dioxide and residual methane, (i.e., synthesis gas or syn-gas). In a second step, 4 the syn-gas is converted into methanol. 5 6 The second step of converting the syn-gas into methanol is a well-known 7 process. It typically involves a catalytic process using a copper-based catalyst, 8 such as a catalyst comprising a reduced zinc oxide/copper oxide mixture, among 9 others. To provide the optimum production of methanol from this reaction, the 10 reaction is typically carried out at pressures within the range of 40-100 bars and at 11 temperatures in excess of 200 degrees C and blow 320 degrees C, with a 12 temperature range of between 220 and 280 degrees C being most common. The 13 production of the syn-gas itself is typically carried out at pressures within the 14 range of 20-40 bars depending on the reformer technology that is utilized. 15 16 Due to the particular mechanism of the reaction for the production of 17 methanol, the reaction does not go to completion, as the concentration of produced 18 methanol is limited by equilibrium. Specifically, the amount of methanol 19 contained in the product gas exiting the reactor comprises about 6-8 mol% of the 20 total gas, although it can be higher. This methanol is removed from this product 21 gas stream by condensing it through cooling the product gas stream to below 110 22 degrees C, and most commonly below 60 degrees C. The cooled methanol can 23 then be removed from the gas stream while the excess syn-gas is sent back to the 24 reactor in order to further react the excess syn-gas. This enables additional 25 methanol to be obtained from the syn-gas recycled back to the reactor in 26 combination with an amount of fresh syn-gas that is also charged to the reactor. 27 28 In perfonning this recycle step, one well-known process involves the use 29 of a recycle compressor which receives the excess syn-gas from the separator and 30 compresses it in order to overcome the pressure drop that occurs within the reactor 31 and separator. This type of reactor is commonly referred to as'a recycle loop 3 WO 20071096699 PCT/IB2006/003279 1 reactor and is schematically shown in Fig. 1. In this reactor, the concentration of 2 methanol in the syn-gas leaving the reactor is low enough that the volumetric flow 3 rate of the excess syn-gas through the recycle compressor is typically two (2) to 4 ten (10) times the volumetric flow rate of the fresh syn-gas being introduced into 5 the reactor separately from the syn-gas charged to the reactor from the recycle 6 compressor. A purge gas stream of approximately 4 to 8 percent of the recycled 7 syn-gas stream is also expelled from the recycle loop, prior to re-compression to 8 control the concentration of inert material that builds up in the reactor as a result 9 of the recycle, 10 11 One significant drawback to the recycle loop reactor described above is the 12 cost of the recycle compressor. Often the recycle function is incorporated into a 13 single drive train compressor that compresses the syn-gas up to the pressure of the 14 recycle loop and also provides for re-compression of the recycle gases. The 15 compression train is an expensive item of equipment and may be the single most 16 expensive purchased component in the construction of a methanol producing 17 facility. As stated previously, a high recycle flow is utilized to enable high overall 18 conversion of syn-gas to be achieved. The recycle compressor also becomes 19 significantly cheaper per unit volume compressed as the scale of the plant is 20 increased. Thus, the use of a process using a recycle compressor is, as a practical 21 matter, preferential for those facilities producing a relatively high daily output of 22 methanol relative to the current maximum single train production plant, such as 23 those facilities producing in the neighborhood of 5,000 tons per day in 2005, 24 where maximum efficiency is required and integration df the syn-gas compressor 25 and recycle compressor can be achieved in order to make the use of a recycle 26 compressor economically viable. 27 28 As an alternative to a recycle loop reactor, it has been conceived that the 29 facility could utilize a multiple reactor set or cascade process (Fig. 2) whereby the 30 syn-gas is initially fed to a first reactor for reaction with the catalyst contained 31 therein to produce methanol. The product gas then enters a first separator or 4 WO 2007/096699 PCT/IB2006/003279 I knock-out pot wherein the methanol produced in the first reactor is cooled into 2 liquid form and separated from the excess syn-gas. The first separator can operate 3 to separate the methanol for the syn-gas in any desired manner, such as by gravity 4 or by applying centrifugal focus to the products. The remaining excess syn-gas is 5 then fed to a second reactor, which undergoes the same reaction, thereby 6 producing additional methanol. The additional methanol is removed from the 7 second reactor and directed to a second separator in the same manner. The 8 number of reactors and separators can be selected to create a multiple reactor set 9 that achieves the desired conversion percentage of the syn-gas to methanol. For 10 example, when using an optimal syn-gas composition that has a conversion to 11 methanol of 50 percent in each reactor, the reactor set can be selected to include 12 four reactors and four separators, which theoretically results in the achievement of 13 a 95 percent overall conversion of the syn-gas to methanol after the fourth 14 separator. 15 16 However, in more realistic situations, when the syn-gas composition is less 17 optimal, such as when the stoichiometry of the syn-gas is away from the 18 stoichiometry of the reaction desired, e.g., the ratio ([moles hydrogen]-[moles 19 carbon dioxide])/([moles carbon monoxide]+[moles carbon dioxide]) is between 20 2.5 and 3.0, then more than four reactor sets and perhaps as many as 10 reactor 21 sets are required for high conversion (>95%). Additionally, when there are 22 stoichiometric quantities of reactant in the syn-gas, but high levels of inerts also 23 present, then a high number of reactor sets will be required, such as where 24 autothermal reforming with air is performed for the production of syn-gas 25 resulting in dilution of the syn-gas with high levels of nitrogen. 26 27 There are several areas in a methanol production process where the ability 28 to employ a cascade of reactors would be considered beneficial. Principally the 29 benefit from using the cascade comes from not requiring a recycle compressor. 30 Furthermore, it should be borne in mind that in a grass roots methanol production 31 plant the recycle compressor is often part of the compressor associated with syn 5 WO 20071096699 PCT/IB2006/003279 1 gas compression. Thus, eliminating the syn-gas compressor along with the 2 recycle compressor in a grass roots installation gives the maximum benefit. 3 4 While the use of a simple cascade of reactors for this particular purpose is 5 disclosed in certain prior art references, the references that discuss the use of such 6 a cascade of reactors focus exclusively on methods by which the reformer 7 operating pressure can be matched to the methanol synthesis pressure. For 8 example, U.S. Patent Nos. 5,177,114; 5,245,110; 5,472,986; and 7,019,039 are 9 each patents that disclose inventions in the field of autothermal reforming using 10 air rather than oxygen. However, while these patents very generally disclose the 11 use of cascade reactors in the methanol production process, they do not address 12 the issues of how the cascade of reactors can be made cost effectively. 13 Furthermore, each of U.S. Patent Nos. 5,177,114; 5,245,110; and 5,472,986 14 disclose a methanol production process where the recycle compressor can be 15 eliminated as a result of operating a reformer autothermally, and then converting 16 the syn-gas to a methoxy compound using three to five reactor sets with product 17 condensation between each stage. Recognizing conventional wisdom that a 18 cascade process cannot achieve a high syn-gas to methanol conversion, carbon 19 efficiencies for the methanol synthesis section of less than 80% are quoted, 20 whereas in conventional plant efficiencies in excess of 95% are achievable. 21 22 Additionally, U.S. Patent No. 6,255,357 discloses a methanol production 23 process that uses pressurization of the oxidant gas for fired heating of the steam 24 reformer as a means of achieving an mechanically feasible high pressure steam 25 reformer with an operating pressure sufficient to ensure a sufficient operating 26 pressure throughout the process. The process also includes a cascade of reactors 27 downstream from the reformer in which the reformed syn-gas is converted to 28 methanol. The pressurization of the incoming natural gas into the reformer avoids 29 the requirement of a syn-gas compressor upstream of the cascade of reactors, as 30 well as avoiding the need for a recycle compressor. However, as with the 6 WO 20071096699 PCT/IB2006/003279 1 previous references, the cascade of reactors is only very generally disclosed 2 without any discussion as to how the cascade can be made economically. 3 4 Other situations regarding a methanol production process where it would 5 be considered advantageous to avoid use of a recycle compressor include those 6 where the compressor would be added to an existing methanol production facility 7 in a retro-fit capacity, or as part of an addition to a planned methanol production 8 facility construction for the purpose of removing any remaining methanol from 9 the purge gas discharged from the facility. One system of this type that addresses 10 the loss of the potential and actual methanol present in the purge gas'stream is 11 disclosed in U.S. Patent No. 6,258,860, which is incorporated by reference herein 12 in its entirety. The process disclosed directs the purge gas stream produced by a 13 methanol synthesis zone to another methanol synthesis or production zone in 14 order to both collect the methanol present in the purge gas stream as well as to 15 further react the unreacted components of the purge gas stream to produce 16 additional methanol. 17 18 However, the process disclosed in the '860 patent has certain drawbacks in 19 that it utilizes a compressor to compress the combined purge gas and recycled 20 syn-gas stream prior to further reacting the combined stream. Because, as 21 discussed previously, the recycle compressor is the highest cost item in a 22 methanol production system, the use of additional recycle compressors to recover 23 methanol from a purge gas is highly undesirable, especially for systems producing 24 a relatively low daily output of methanol relative to the current maximum single 25 train production plant, such as those facilities producing in the neighborhood of 26 5,000 tons per day in 2005. 27 28 Another example where additional compressor capacity may be avoided 29 through the use of a set of cascade reactors could occur as part of a re-vamp or de 30 bottle-necking of a methanol plant. If the re-vamp or de-bottlenecking entails 31 increased syn-gas production, then the capacity of the methanol converter would 7 WO 2007/096699 PCT/IB2006/003279 1 be required to be increased. It may be possible to increase the effectiveness of the 2 reactor through better packing of catalyst or dividing the catalyst into multiple 3 beds in a single reactor. However, where the reactor already makes effective use 4 of the catalyst, it may not be possible to economically increase the performance of 5 the reactor. Further, another limitation on the operation of the reactor in this 6 situation is the pressure drop of the process gas across the reactor. Increasing the 7 re-circulation rate, increasing the catalyst volume, increasing the feed flow or 8 reducing the purge rate will all increase the pressure drop of the process gas 9 through the reactor. There will of course be a consequent limitation on the 10 capacity of the re-circulation compressor as well to recompress the gas for 11 reintroduction into the reactor, 12 13 One alternative to mitigate these issues would be to operate the methanol 14 converter at reduced conversion conditions, but with a higher gas feed rate, 15 thereby allowing pressure drop limitations to be avoided and then utilizing a 16 separate cascade reactor system to convert the un-reacted gases to methanol 17 without the requirement of an additional compressor or replacing the original 18 reactor. This also has the advantage of being a lower risk method of increased 19 throughput, as the original reactor performance is well known. There will of 20 course be many other circumstances under which a cascade system can be 21 utilized, but all of these circumstances will be reliant on a cost-effective design of 22 cascade system. 23 24 One significant drawback with the multiple or cascade reactor set types 25 utilized in methanol production as described in the prior art results from the end 26 use construction of each reactor, heat exchanger(s) and separator forming the 27 individual reactor set. Specifically, because a portion of the syn-gas is lost in each 28 reactor set based on its conversion to methanol, often each subsequent reactor set 29 and separator is constructed to be smaller than the immediately preceding ones to 30 accommodate the reduction in the flow rate of incoming syn-gas. This might 31 initially be anticipated to be highly beneficial based upon the reduction in the 8 WO 2007/096699 PCT/IB2006/003279 1 amount of material necessary to construct each successive reactor set. However, 2 each reactor set requires the same functionality, connections, cooling and access 3 for catalyst replacement, which become more difficult and/or expensive to 4 manufacture on a progressively smaller scale. In addition, the cooling, gas-liquid 5 separation and re-heating of the methanol-bearing stream as it passes between the 6 various reactor sets must be effected in an energy efficient and cost effective 7 manner. Further, all of the reactor sets and separators must be constructed to be 8 operable at the elevated pressures (40-100 bars) that the reactions occurring for 9 the conversion of the syn-gas to methanol require. 10 11 One example of a system that attempts to address this shortcoming is 12 disclosed in U.S. Patent No. 6,723,886 in a methanol production process using 13 reactive distillation. However, while there is removal of methanol between 14 reactor beds by condensation within the reactor, the condensation takes 15 necessarily takes place at reaction temperature, and condensation at elevated 16 temperature limits the conversion of methanol to approximately 60%. However, 17 even with the significant restriction this places on methanol production, this is in 18 accord with the current industry view that condensation at reduced temperature is 19 not viable. 20 21 Therefore, it is desirable to develop a multiple or cascade reactor set and a 22 process for the production of products of equilibrium limited reactions, e.g., 23 methanol, using the multiple reactor set to obtain a high percentage conversion of 24 feed syn-gas to methanol by condensing the methanol in the reactor effluent in an 25 interstage feed/effluent heat exchanger. It is also desirable that the multiple 26 reactor set be operable without the need for a gas recycle compressor and 27 preferably without the need for the construction of multiple individual reactors, 28 heat exchangers, and separators. In other terms, the heat exchanger design should 29 be suitable for efficient operation and integration into the reactor sets, while also 30 minimizing the number of necessary equipment items. 31 9 WO 20071096699 PCT/IB2006/003279 1 With regard to the goal of minimizing the necessary number of equipment 2 items in a reactor set, it is easier to understand the conventional approach to 3 solving this problem of eliminating equipment items and reducing the cost of 4 equipment items by reference to the specific problems of a conventional methanol 5 synthesis loop. Apart from the recycle compressor, a methanol synthesis loop 6 contains six principle operations: 1) pre-heat of the gas; 2) reaction of the gas to 7 fonn methanol; 3) removal of the heat of reaction as high grade heat; 4) cooling of 8 the gas to methanol condensation temperatures; 5) condensation of the methanol 9 using cooling water; and 6) vapor/liquid separation. In a typical plant there may 10 be two integrations of these functions for the purposes of minimizing the 11 necessary equipment items which are removal of the heat of reaction is performed 12 by steam raising in a shell and tube reactor, and pre-heat of the gas by feed 13 effluent exchange. Thus, a typical synthesis loop will consist of at least six 14 equipment items: 1) a start up heater; 2) a feed/effluent heat exchanger; 3) a 15 reactor; 4) a high grade heat recovery unit; 5) a water cooler; and 6) a gas-liquid 16 separator. 17 18 Steam raising directly in the reactor does eliminate the requirement for a 19 separate high-grade heat recovery unit. However, it also requires a steam drum 20 with the reactor and so does not reduce the number of equipment items. 21 22 With regard to the use of the feed/effluent heat exchanger, the highest 23 energy efficiency is achieved with a high effectiveness heat exchanger that is able 24 to maximize the cooling of the effluent stream. Increasing the amount of high 25 grade heat recovered reduces the temperature difference in the feed/effluent 26 exchanger. Therefore, for maximum high grade heat recovery a high 27 effectiveness heat exchanger is required. However, shell and tube heat 28 exchangers as used in prior art multiple reactor sets can only achieve high 29 effectiveness through the coupling of multiple heat exchanger units, again 30 increasing the number of equipment items required. The usefulness of high grade 31 heat recovery is, in part dependent, on the temperature at which it is recovered. In 10 WO 2007/096699 PCT/IB2006/003279 1 particular, for a methanol process the high grade heat recovery from the methanol 2 synthesis section is utilized for steam raising for the refonner. This requires that 3 the stream from which heat is being recovered is above a minimum temperature, 4 typically 200-250 deg C. However methanol condensation temperatures are in the 5 region of 60-100 deg C. For efficient operation, therefore, heat exchangers are 6 required that can operate with a hot gas temperature span of approximately 150 7 deg C. The heat of reaction is recovered by cooling the reactant stream by the 8 equivalent of typically 50-100 deg C of sensible heating. If the feed gas is 9 introduced to the methanol reactor at a temperature below high grade heat 10 recovery temperature this represents a loss of energy efficiency in the system and 11 increases the low grade cooling requirement. Consequently the temperature 12 difference in the feed-effluent exchanger will be kept to less than 50 deg C and 13 typically 20-30 deg C. Where high single pass conversions can be achieved in the 14 reactor, such as with a balanced stoichiometry, high operating pressure, efficient 15 heat removal or a low overall conversion the temperature constraints may be more 16 relaxed. However, this often brings greater reactor complexity or lower overall 17 efficiency. 18 19 The performance measure of a heat exchanger can be described in terms of 20 temperature span and log mean temperature difference between stream. The value 21 (span divided my lmtd) is referred to as NTU count and as can be seen above it 22 would be desired for an energy efficient methanol process that the fee/effluent 23 exchangers would operate with an NTU count above 5, and more preferably above 24 7. 25 26 The problem concerning the number of equipment items is also not 27 alleviated when a recycle loop is replaced with a cascade system. With no recycle 28 there is no requirement for a recycle compressor. However, for each contact with 29 the catalyst there will be up to six additional equipment items, as discussed 30 previously. One option to reduce the number of equipment items is to eliminate 31 some of the heat exchangers. For example, instead of recovering high grade heat 11 WO 2007/096699 PCT/IB2006/003279 1 from the reactor gases, the gases can be used to directly heat the incoming feed 2 gases. The feed/effluent exchanger is then smaller as a result of an increased 3 driving temperature, but the reaction heat is thcn lost to the cooling water and a 4 less efficient process is produced. 5 6 Therefore, to improve the economics and efficiency of the prior art 7 methanol cascade systems, it is necessary to solve the following issues: 1) to 8 minimize the number of equipment items; 2) to increase the effectiveness of the 9 feed/effluent exchangers; and 3) to integrate multiple functions into single 10 equipment items. 11 12 SUMMARY OF THE INVENTION 13 One method by which the improvement of the economics and efficiency of 14 the cascade reactor system can be achieved, and which is an integral part of the 15 apparatus and method of the present invention, is the use of extended surface, or 16 plate style heat exchangers in the cascade reactor system. In particular, plate-fm 17 (brazed or diffusion bonded) or printed circuit heat exchangers (PCHE) are able to 18 achieve high effectiveness eliminating the requirement for multiple units for a 19 single duty or a pre-heater to reduce the required effectiveness. The construction 20 methods of plate style heat exchangers also allow for multiple stream heat 21 exchangers to be combined into a single unit. For example, the reactor outlet 22 gases can pass through a single plate style heat exchanger where, in a first section, 23 the gas is cooled with a high temperature coolant such as pressurized water at 200 24 250 deg C. In a second section of the heat exchanger, the reactor gases are cooled 25 by thermal contact with the reactor inlet stream. Finally, in a third section, the 26 reactor gases are cooled with cooling water to condense the water. 27 28 When a series or cascade of reactors is used, as discussed previously, the 29 appropriate design for heat recovery may differ as the reactor diminish in size. In 30 the early reactors of the set the cost of heat recovery is more economical as the 31 energy recovered per unit is greatest. With each successively smaller reactor and 12 WO 20071096699 PCT/IB2006/003279 1 associated heat exchanger, the amount of energy that is available per unit 2 diminishes as the rate of production of methanol in the set is smaller and the costs 3 of heat recovery can become prohibitive. So, the present invention can utilize as 4 the first reactor a steam raising or gas tube cooled reactor within a recycle loop, as 5 discussed previously. The subsequent reactor sets each utilize a high efficiency 6 heat exchanger of the aforementioned type integrating high grade heat recovery, 7 feed effluent heat exchange and methanol condensation utilizing cooling water, 8 whereas the final reactor set only utilizes a high efficiency heat exchanger for 9 feed/effluent heat exchange and water cooling. 10 11 For even smaller methanol production units, it may also be possible to 12 incorporate the heat exchange functions, i.e., the high efficiency heat exchangers, 13 for each reactor set into a single fabricated unit. This is made possible by the use 14 of plate style heat exchangers that are amenable to such construction. 15 Additionally, separate from the integration of the heat exchangers into a single 16 unit, integration of the reactors themselves and also the vapor/liquid separators 17 into the same or separate units are also possible. At this smaller scale of 18 production, the present invention can have all of the reactors contained within a 19 single item of equipment, and all of the heat exchangers in a single unit, as well as 20 all of the vapor/liquid separators. As a result, the cascade methanol process is 21 effectively reduced to only three principle equipment items. 22 23 The integration of these units allows for several different uses of the 24 integrated units. For example, a cascade reactor set unit of 3 or 4 reactors could 25 be connected to the purge gas stream from a methanol synthesis loop to allow for 26 the further reaction of the contents of the purge gas stream through the unit to 27 form additional methanol. This would increase the overall conversion percentage 28 of the loop without increasing the recycle rate, as only the purge gas is directed 29 through the cascade reactor set unit. Also, the addition of the cascade reactor unit 30 would both increase the amount of methanol that can be made from a fixed stream 31 of methane, or, as part of a wider retro-fit to an existing production process, would 13 WO 20071096699 PCT/IB2006/003279 1 boost the capacity of the methanol synthesis section without an increased gas rate 2 through the recycle compressor. 3 4 Therefore, according to a first aspect of the present invention, an improved 5 multiple or cascade reactor set type chemical production system is provided in 6 which the incoming reaction gas entering the first reactor is an amount of excess 7 reaction gas and/or a purge gas from a conventional single reactor, a cascade 8 reactor set or a recycle loop reactor system, each of which are fed from a steam 9 reformer or an autothennal reformer. The gas initially passes through a first high 10 efficiency plate-type or extended surface heat exchanger, whereby the incoming 11 syn-gas comes into thermal contact with some or all the gaseous reaction products 12 exiting the first reactor in order to cool the gaseous reaction products and preheat 13 the incoming syn-gas. Some or all of the gases exiting the first reactor may be 14 brought into thermal contact with an additional stream for the recovery of high 15 grade heat prior to entering the feed-effluent heat exchange section. The cooling 16 of the reaction products causes the desired product to condense into liquid form 17 within the heat exchanger. Further cooling of the product bearing stream is then 18 effected by thermal contact with an additional cooling utility stream that is also 19 introduced into the aforementioned heat exchanger, such an arrangement being 20 known as a multi-stream heat exchanger. The cooled reaction products, including 21 the condensed liquid product, can be removed directly from the heat exchanger or 22 can then flow into the first separator where the condensed liquid product is 23 removed in the separator while the excess reaction gas is directed into a second 24 reactor for additional production of the desired chemical. Prior to reaching the 25 second reactor, the excess reaction gas passes through a second high-efficiency 26 heat exchanger to be warmed by the gaseous reaction products exiting the second 27 reactor, and consequently condenses the product contained in the product gases 28 exiting the second reactor. The high-efficiency plate-type or extended surface 29 heat exchangers condense the desired product, for example methanol, produced in 30 each of the reaction zones in a highly economical manner, as each high-efficiency 31 heat exchanger has a close temperature of approach with a counter-current design 14 WO 20071096699 PCT/IB2006/003279 1 to minimize the amount of cooling water necessary and to maximize the amount 2 of steam recovery. Further condensation is achieved through the introduction of 3 an additional cooling stream to the high-efficiency heat exchanger so as to effect 4 thermal contact with the product bearing stream and increase the amount of 5 product condensed from the product bearing stream. Further, contrary to other 6 references which state that the interstage removal of methanol by condensation is 7 not practical or economical, such as K. R. Westerterp, New Methanol Processes, 8 "Energy Efficiency in Process Technology" Ed. P. A. Pilarvachi, Elsevier Applied 9 Science, 1993, pp. 1142-1153, 1147, the use of the high-efficiency plate type or 10 extended surface heat exchangers to provide interstage methanol condensation 11 operates in both a practical and economically viable manner within the apparatus 12 and method of this invention. 13 14 According to another aspect of the present invention, the various reactor 15 sets of the cascade system are formed as reaction zones integrated within a single 16 reactor vessel, such that only the reactor vessel and appropriate inlet and outlet 17 fittings on the vessel need to be constructed to withstand the temperatures and 18 pressures necessary for the methanol production reaction. The reaction products 19 from each reaction zone are passed through the high-efficiency heat exchangers 20 which are also formed in a block-like configuration positioned and connected 21 between each reaction zone in the reactor vessel, and a separator zone located in a 22 separate separator vessel constructed similarly to the reactor vessel. The 23 construction of the various reaction zones within the reactor vessel and the various 24 separator zones within the separation vessel greatly reduces the cost of the 25 materials necessary to construct the various vessels, as the pressure differentials 26 between the respective zones in each of the reactor vessel and the separation 27 vessel are minimal. This eliminates the need for constructing individual walls 28 between the various zones of materials capable of withstanding the high pressure 29 differentials between reaction pressure and atmospheric pressure that would 30 otherwise be encountered. 31 15 WO 2007/096699 PCT/IB2006/003279 I According to still a further aspect of the present invention, an improved 2 multiple or cascade reactor set type methanol production system is provided in 3 which methanol is initially produced through the use of any suitable methanol 4 production system, such as a conventional methanol synthesis loop with a recycle 5 compressor. The purge gas stream from the methanol production system is 6 subsequently directed through a cascade system of three or more reactor sets 7 formed according to the present invention that further react the unreacted 8 components of the purge gas stream to form additional methanol. The reactors 9 are constructed separately, or as part of a single vessel with a separating wall 10 designed to contain the differential pressure between the reactors and provide II access between reactors to aid filling of the individual beds with catalyst. In each 12 reactor set a single multi-stream heat exchanger is used to recover high grade heat, 13 to effect feed/effluent heat exchange, and to achieve condensation of methanol by 14 further cooling with a cooling medium such as water. In addition, the compact 15 multi-stream heat exchangers are arranged in such a manner alongside the reactor 16 set so as to minimize the amount of connections between the cascade reactor set 17 and the synthesis loop. 18 19 According to still another aspect of the present invention, the multiple 20 reactor set utilizing the high-efficiency plate-type or extended surface heat 21 exchangers can be utilized as a stand-alone stationary or mobile system and/or as 22 an add-on to an existing recycle loop reactor set or to an existing cascade reactor 23 set to further increase the percent conversion of methanol from these pre-existing 24 reactors, or maintain the overall conversion of the modified process while relaxing 25 the effectiveness of the recycle process through, for example, a reduced recycle 26 rate. 27 28 Numerous additional aspects, features, and advantages of the present 29 invention will be made apparent from the following detailed description together 30 with the drawing figures. 31 16 WO 20071096699 PCT/IB2006/003279 1 BRIEF DESCRIPTION OF THE DRAWINGS 2 3 The drawings illustrate the best mode currently contemplated of practicing 4 the present invention. 5 6 In the drawings: 7 Fig. 1 is a schematic view of a prior art recycle loop reactor system; 8 Fig. 2 is a schematic view of a prior art multiple reactor set or cascade 9 system; 10 Fig. 3 is a schematic view of the multiple reactor set of Fig. 2 including a 11 number of interstage feed/effluent heat exchangers constructed according to the 12 present invention; 13 Fig. 4 is schematic view of a second embodiment of the reactor set of Fig. 14 3, in which the separate reactors, heat exchangers and separators are formed as 15 zones disposed within a single vessel; 16 Figs. 5A-5C are isometric views of header constructions for the 17 feed/effluent heat exchangers of the reactor of Fig. 4; 18 Fig. 6 is a schematic view of a weir discharge system for the separators of 19 the reactor of Fig. 4; and 20 Fig. 7 is a schematic view of the reactor set of Fig. 3 attached to the purge 21 gas stream of a methanol production reactor system. 22 23 DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS 24 25 With reference now to the drawing figures in which like reference 26 numerals designate like parts throughout the disclosure, a methanol production 27 system that can be utilized with or replaced by the present invention is indicated 28 generally at 100 in Fig. 1. The system 100 shown is a conventional methanol 29 synthesis loop, that receives a stream 160 of syn-gas from a reformer 120, 30 converts a portion of the syn-gas to methanol in a reactor 140, and discharges a 31 combined stream 180 of unreacted syn-gas and methanol to a condenser or 17 WO 2007/096699 PCT/IB2006/003279 I separator 200. The combined stream 180 is separated in the separator 200 into a 2 methanol stream 220 that is collected, and a recycle stream 240 that is directed 3 from the separator 200 to a recycle compressor 260. The compressor 260 4 compresses the gas in the recycle stream 240 and directs the recompressed gas 5 stream 280 back to the reactor 140. However, a portion of the recycle stream 240 6 entering the compressor 280 is diverted as a purge gas stream 300 prior to 7 recompression. 8 9 In addition, another prior art methanol production system that is capable of 10 use with the present invention is the multiple reactor set or cascade system 320 11 shown in Fig. 2. Similar to the recycle system 100, the cascade system 320 also 12 includes the reformer 120 that forns and directs a stream 160 of syn-gas to a first 13 reactor 140 that converts a portion of the stream 160 to methanol, forming a 14 combined stream 180 that exits the reactor 140. This combined stream 180 15 subsequently enters a first separator 200 where it is separated into a methanol 16 stream 220 and a first stream 340 of excess syn-gas. This excess syn-gas stream 17 340 is then directed to a second reactor 360 for further reaction of the stream into t 8 methanol. The resulting second combined stream 380 is then passed to a second 19 separator 400, where a second methanol stream 220' is fonned and combined with 20 stream 220, and a second excess syn-gas stream 340' is passed to a third reactor 21 360'. The third reactor 360' forms from the second excess stream 340' a third 22 combined stream 380' that passes into a third separator 400' and which exits as a 23 third methanol stream 220" and a third excess syn-gas stream 340". The third 24 methanol stream 220" is combined with streams 220 and 220', and the third 25 excess stream 340" is passed to a fourth reactor 360". The reactor 360" 26 transforms the excess stream 340" into a fourth combined stream 380", that enters 27 a fourth separator 400". The fourth separator 400" separates the combined stream 28 380" into a fourth methanol stream 220"' and a purge gas stream 420. 29 30 A multiple reactor set methanol production system constructed according 31 to the present invention is indicated generally at 10 in Fig. 3. The system 10 18 WO 2007/096699 PCT/IB2006/003279 1 receives a stream 16 of syn-gas from a reformer (not shown), converts the syn-gas 2 to methanol in stages, and ultimately discharges a stream 21 of purge gas and a 3 combined methanol stream 25. Significantly, it lacks a recycle compressor yet 4 still operates efficiently, permitting it to be used economically in relatively small 5 scale applications producing less than 2,500 tons of methanol per day, and more 6 particularly on the order of less than 1,500 or less than 1,000 tons of methanol per 7 day. It could, however, be scaled up for larger scale operation as well or reduced 8 in size without significant detriment. 9 10 The syn-gas typically contains approximately 66mol% hydrogen, 20 11 mol% carbon monoxide, 9 mol% carbon dioxide, and 2mol% methane. It would 12 also contain any nitrogen that was present in the methane originally fed to the 13 reformer where the syn-gas is made. The actual composition will depend on the 14 pressure and temperature used in the reforming, the method of reforming (steam 15 reforming, autothermal etc.) and whether there was any carbon dioxide added to, 16 or present in the methane stream fed to the refonner. 17 18 The system 10 includes a number of reactor sets 11, 13, 15, 17 located in 19 series such that each downstream reactor set receives the effluent from the 20 immediately upstream reactor set as a feed stream, further conveys the feed 21 stream, and discharges a condensed methanol stream 24 and an effluent stream 20. 22 The methanol streams 24 are combined to form combined stream 25. The effluent 23 stream from the downstream-most reactor set forms the purge gas stream 21. 24 25 Still referring to Figure 3, each of the reactor sets 11, 13, 15, 17 includes a 26 reactor 12, a separator 18, and a feed/effluent heat exchanger 22. Each of the 27 reactors 12 receives a feed stream feed 16 or 20 and discharges a product stream 28 14. The product stream 14 from each of the reactors 12 is directed through a 29 corresponding heat exchanger 22, where it is cooled by heat exchange from the 30 feed stream 16 or 20 for the reactor of that set to condense the methanol from the 31 product stream. The stream 14 is additionally cooled by the introduction of 19 WO 20071096699 PCT/IB2006/003279 1 cooling utility within the structure of the feed-effluent exchanger 22 such that 2 thermal contact is achieved. The product stream 14 of each reactor set 11, 13, 15, 3 17 is then directed to a corresponding separator 18 of any suitable type which, in 4 turn, provides a return stream 20 to the reactor 12 of the next downstream reactor 5 set. As mentioned above, the return stream 20 from the separator 18 of the final 6 reactor set 15 is discharged as a purge gas stream 21. 7 8 The reactors 12 can be selected to be any suitable type of reactor for use in 9 a methanol production reaction as are known in the art. Preferably each reactor 10 consists of multiple adiabatic beds with cooling of the process fluid after each bed 11 such that steam is produced to be used as a utility. Only two adiabatic beds per 12 reactor are sufficient to enable the correct operation of the process. Suitable 13 alternate reactors could be conceived by comparison with conventional methanol 14 reactors such as steam jacketed tubes (or Lurgi reactor), a tube cooled reactor 15 available from Davy-Synetix, an adiabatic bed reactor with more than two beds, or 16 perhaps spherical or radial geometry multiple adiabatic bed reactors, among 17 others. Each reactor 12 contains a methanol synthesis catalyst such as one 18 comprising a reduced zinc oxide/copper oxide mixture. The conversion typically 19 takes place at 40-100 bars and temperatures in excess of 200'C, typically at 220 20 280'C but not exceeding 310'C. Typically the product stream 14 will contain 21 approximately 5mol% methanol. Higher conversion rates can be achieved at 22 lower temperatures but at the expense of larger catalyst volumes. 23 24 The heat exchangers 22 are preferably selected to be plate-type heat 25 exchangers, such as a diffusion bonded printed circuit heat exchanger 26 manufactured by Heatric of Dorset, UK, or extended surface heat exchangers, 27 such as a diffusion bonded plate fin heat exchanger available from Heatric, or a 28 brazed plate fin heat exchanger available from Chart Industries of Bracknell, UK, 29 a spiral wound heat exchanger, or other suitable stacked plate heat exchangers, as 30 opposed to prior art shell and tube heat exchangers. These types of heat 31 exchangers are preferred because plate-type or extended surface heat exchangers 20 WO 2007/096699 PCT/IB2006/003279 1 22 are capable of providing a close temperature of approach to the fluid streams 2 passing through the exchangers 22, such that the product streams exit the 3 exchangers 22 within five (5) degrees C of one another. This effectively 4 minimizes the amount of cooling water required and maximizes the steam 5 recovery from the heat exchangers 22, such that these types of heat exchangers 6 have an effectiveness greater than 7 NTU. 7 8 The heat exchangers 22 also are capable of multi-stream and/or 9 countercurrent operation such that multiple heat transfer operations can take place 10 within a single heat exchanger 22. Specifically, the heat exchangers 22 effectively 11 cool the methanol component within each of the product streams 14 to condense 12 the methanol in the product streams 14 and enable removal of the methanol within 13 the separators 18 in a highly economic manner. An additional cold utility stream 14 is introduced into the heat exchanger to maximize methanol condensation. In 15 addition, the heat retained by the product streams 14 is effectively utilized to 16 elevate the temperature of the feed stream 16 and the return streams 20 prior to 17 these streams 16 and 20 entering the reactors 12. Methanol typically begins to 18 condense at 110'C, depending on concentration and reaction pressure. For 19 efficient removal of methanol (>75%) where the exit concentration is typical 20 (5%), then the heat exchangers 22 are preferably configured to cool the product 21 streams 16 to below 60'C by the incorporation of a cooling utility stream into the 22 feed/effluent exchange reactor. Most preferably, the reactor return stream 20 can 23 pass through a single plate style heat exchanger 22 where, in a first section, the 24 gas is cooled with a high temperature coolant such as pressurized water at 200 25 250 deg C. In a second section of the heat exchanger 22, the reactor gases are 26 then cooled by thermal contact with the reactor inlet stream 16. Finally, in a third 27 section, the reactor gases are cooled with cooling water to condense the water. 28 29 While the system 10 shown in Fig. 3 illustrates four reactors sets 11, 13, 30 15, 17, the number of reactor sets and constituent components of each reactor set 31 can be varied as desired. For instance, optimal operation of the illustrated 21 WO 2007/096699 PCT/IB2006/003279 1 embodiment with four reactors 12 requires a tight composition control of the syn 2 gas in the feed stream 16 in order to keep the stoichiometric ratio [(H2 3 C02)/(CO+CO2)] between 2:1 and 3:1, and preferably between 2.1-2.2/1 in order 4 to achieve the 95 percent conversion of the syn-gas to methanol required for an 5 economically viable process. However, by adding additional reactor sets, up to 6 ten (10), a system would capable of producing acceptable overall syn-gas 7 conversions (i.e., in excess of 95 percent for COx), or in excess of 90 percent for 8 H2) with a wider range of feed gas compositions having ratios between 3:1 and 9 2:1, and/or for feed gases with changing compositions, such as are present at C02 10 stimulated gas fields. 11 12 Referring now to Fig. 4, in a more specific embodiment of the invention, 13 the system 10' includes a reactor vessel 27 that defines a number of reaction zones 14 26 therein, each of which contains the reactor 12 for a corresponding reactor set 15 11', 13', 15', etc. The system 10' also includes a separator vessel 30 which 16 defines a number of separation zones 32 therein for each of the reactor sets 11', 17 13', 15', etc. The vessels 27 and 30 are constructed in a manner which allows the 18 vessels 27 and 30 to withstand the elevated temperature (200'C to 310'C) and 19 pressure (40-100 bars) required for the methanol production reaction. To separate 20 the various reaction zones 26, dividing walls 28 are disposed between adjacent 21 zones 26. Because all of the reaction zones 26 are disposed within the vessel 27, 22 the only pressure differential between the zones 26 is the pressure drop between 23 the process streams flowing into and out of successive reaction zones 26, which is 24 typically around 0.2-2.0 bars. Thus, the dividing walls 28 are constructed of 25 materials that only need to be able to withstand a pressure differential of around 26 0.2-2.0 bars, which is much less costly than the materials forming the vessel 27, 27 and the walls 28 can be of a simple welded construction. Ease of construction 28 may result in a design that uses more than one reactor vessel to accommodate the 29 multiple reactor zone and more than one separation vessel to accommodate the 30 methanol separators. Additionally, access to each of the zones 26 can be provided 31 through the walls 28 at an internal access point (not shown) capable of 22 WO 2007/096699 PCT/IB2006/003279 1 withstanding 2 bars of pressure, instead of an external access point (not shown) 2 requiring an forty (40) to one hundred (100) bar design pressure. Also, with the 3 inclusion of the reactors 12, heat exchangers 22 and separators 18 in the various 4 vessels 27 and 30, and heat exchanger block 40, to be discussed, the system 10 5 can be operated at the same pressure as the reformer (not shown) that supplies the 6 syn-gas to the system 10, thus eliminating the need for syn-gas compression prior 7 to being fed to the system 10. 8 9 Similarly, the specific embodiment of the system 10' also includes a 10 separation vessel 30 that defines a number of separation zones 32 or knock out 11 pots, each of which contains the heat exchanger 22 and separator 18 for a 12 corresponding reactor set 11', 13', 15'. The separation zones 32 are separated 13 from one another by dividing walls 34. Again, because the pressure drop between 14 adjacent separation zones 32 is very low, e.g., less than two (2) bars, the dividing 15 walls 34 can be constructed of materials similar to walls 28 and much less costly 16 than the materials utilized for the construction of the separation vessel 30. 17 Additionally, as shown in Fig. 6, due to the small pressure drop between zones 32, 18 a weir system 36 can be utilized that connects each of the zones 32 and allows the 19 condensed methanol stream 24 to flow from a separation zone 32 downwardly 20 along a liquid drain 38 into an adjacent separation zone 32 under the influence of 21 the pressure differentials between the separation zones 32. The liquid drain 38 22 and weir system 36 thereby allow the methanol stream 24 to flow between the 23 respective separation zones 32 or knock-out pots to form the combined stream 25 24 without the need for an active level control of the methanol level within the 25 separation zones 32 and still maintaining a gas seal. In its simplest construction, 26 the zones 32 can be formed with a mixed phase inlet (the product stream 14), an 27 upper gas outlet (the return stream 20), and a lower fluid outlet (the methanol 28 stream 24) when the liquid can directly settle out of the gas in the mixed phase by 29 gravity. 30 23 WO 20071096699 PCT/IB2006/003279 1 Looking at Figs. 4 and 5, the heat exchangers 22 can also be incorporated 2 into a single block 40 where each of the exchangers 22 are positioned to align the 3 entry and exit points (not shown) for the cooling fluid flowing through the 4 exchangers 22 such that they can be connected to a header 42 that, in turn, is 5 operably connected to piping 44 to distribute a cooling fluid flow into the inlet 45 6 for each of the respective heat exchangers 22. As the heat exchangers 22 are 7 formed of extended surface or plate-type heat exchangers, such as printed circuit 8 heat exchangers, a single header 42 can be utilized for a single fluid supplied to 9 each of the heat exchangers 22 in order to provide the cooling water and/or heated 10 steam to the exchangers 22. A similar type of header (not shown) but with 11 dividing walls (not shown) can also be disposed on each exchanger 22 in the block 12 40 that is configured to function as a reaction zone that replaces the zones 26 in 13 the vessel 27 by placing a reaction catalyst in the header, which receives the feed 14 stream 14 or one of the return streams 20 from the separation zones 32. The 15 header allows the methanol conversion reaction to take place within the header 16 and subsequently redirects the product stream 16 with the methanol and excess 17 syn-gas back into the heat exchanger 22 to which the header is connected. 18 19 In the particular embodiments in Figs. 4-6 where the methanol production 20 system 10' includes a reactor vessel 27 with multiple reaction zones 26, a block 21 40 with heat exchangers 22 and a separation vessel 30 with separation zones 32, 22 the system 10' can be configured to be constructed either as a mobile unit, or a 23 fixed unit that has the capacity to synthesize from 1 to 500 metric tons per day of 24 methanol. Also, in addition to the use of the system 10' as a stand-alone unit, the 25 system 10' can be connected to the purge gas stream of a recycle loop reactor or 26 any other methanol production system to increase the conversion of the recycle 27 loop reactor or multiple reactor set by using the purge gas as the feed stream 16 28 for the reactor set 10'. This use is especially advantageous where the reactor set 29 system 10' is formed with the reactor vessel 27, heat exchanger block 40, and the 30 separation vessel 30. 31 24 WO 20071096699 PCT/IB2006/003279 1 Looking now at Fig. 7, another embodiment of the present invention is 2 illustrated in which the system 10 is utilized as a methanol recovery system 52 3 that is operably connected to the purge gas stream 50 of a recycle compressor 260 4 of a recycle loop system 100, similarly to that described previously for the system 5 10'. The system 52 includes a number of reactors 54, 54', 54" operably connected 6 to one another, and which preferably are formed as simple adiabatic reactors. The 7 purge gas stream 50 exiting the recycle compressor 260 initially passes through 8 the a first heat exchanger 56 to preheat the purge gas stream 50 prior to entering 9 the first reactor 54. The purge gas stream 50 is heated by a first reactor product 10 stream 58 that exits the first reactor 54 and passes through the first heat exchanger 11 56 to thennally contact and raise the temperature of the purge gas stream 50. 12 Simultaneously, the purge gas stream 50 lowers the temperature of the first 13 product stream 58 which consists of methanol and still unreacted purge gas. This 14 now-cooled first product stream 58 then passes from the first heat exchanger 56 to 15 a first separator 60 whereby the product stream 58 is further cooled to produce a 16 first methanol stream 62 and a first unreacted purge gas stream 64. The first 17 methanol stream 62 is collected from the first separator 60 to form a methanol 18 product stream 90, while the first unreacted purge gas stream 64 is directed to a 19 second heat exchanger 56' in order to cool and be heated by a second product 20 stream 58' coming from the second reactor 54' of the methanol recovery system 21 52 in the same manner as described regarding first heat exchanger 56. 22 Additionally, the second product stream 58' exiting the second reactor 54' is 23 processed by the second heat exchanger 56' and a second separator 60' in a 24 manner similar to the product stream 58 exiting the first reactor 54 in order to 25 generate a second methanol stream 62' that is collected from the second separator 26 60' and added to the methanol product stream 90, and a second unreacted purge 27 gas stream 64'. 28 29 The second unreacted purge gas stream 64' is directed from the second 30 separator 60' to a third reactor 54" through a third heat exchanger 56" in the same 31 manner as described previously regarding the passage of purge gas stream 50 and 25 WO 20071096699 PCT/IB2006/003279 1 first unreacted purge gas stream 54 through heat exchangers 56 and 56'. The third 2 reactor 54" uses the second unreacted purge gas stream 64" to generate a third 3 product stream 58" that is directed through the heat exchanger 56" to a third 4 separator 60" that generates a collectible third methanol stream 62" that is added 5 to methanol production stream 90, and a third purge gas stream 54" which is 6 discharged from the system 52. 7 8 The number of reactors 54, 54' and 54" can be varied as necessary from 9 one to any number required for the desired methanol conversion, and can be 10 selected to be any suitable type of reactor for use in a methanol production 11 reaction as are known in the art. Preferably each reactor consists of a simple 12 adiabatic reactor, and most preferably with multiple adiabatic beds, with cooling 13 of the process fluid after each bed such that steam is produced to be used as a 14 utility. Only two adiabatic beds per reactor are sufficient to enable the correct 15 operation of the process. Suitable alternate reactors could be selected to be similar 16 to those described previously as alternatives for the reactors 12 utilized in the 17 system 10. 18 19 The heat exchangers 56, 56' and 56" are constructed similarly to the heat 20 exchangers 22 discussed previously, and are preferably selected to be plate-type 21 heat exchangers, such as a diffusion bonded printed circuit heat exchanger, or 22 extended surface heat exchangers, such as a diffusion bonded plate fin heat 23 exchanger or a brazed plate fm heat exchanger, as opposed to prior art shell and 24 tube heat exchangers. These types of heat exchangers are preferred for the same 25 reasons described with regard to the heat exchangers 22 utilized in the system 10, 26 namely due to their ability to provide a close temperature of approach to the fluid 27 streams passing through the exchangers 56, 56' and 56", and the ability of the 28 exchangers to function in a multi-stream and/or counter current manner. This 29 effectively minimizes the amount of cooling water required and maximizes the 30 steam recovery from the heat exchangers 56, 56' and 56". As a result, the heat 31 exchangers 56, 56' and 56" also effectively cool the methanol component within 26 WO 20071096699 PCT/IB2006/003279 1 each of the product streams 58, 58' and 58" to condense the methanol in the 2 product streams 58, 58' and 58" and enable removal of the methanol within the 3 separators 60, 60' and 60" in a highly economic manner. 4 5 Further, the reactor product streams 58, 58' and 58" can be cooled by a 6 utility stream (not shown) prior to entering the heat exchangers 56, 56' and 56". 7 Significantly, it lacks a recycle compressor yet still operates efficiently, permitting 8 it to be used economically in relatively small-scale applications producing on the 9 order of 1,000 to 1,500 tons of methanol per day. It could, however, be scaled up 10 for larger scale operation as well or reduced in size without significant detriment. 11 12 Additionally, the methanol recovery system 52 can include combination 13 elements (not shown) which function as both each of the heat exchangers 56, 56' 14 and 56" and separators 60, 60' and 60" to minimize the number of components 15 utilized in the methanol recovery system 52. 16 17 Further, the methanol recovery system 52 can be utilized as a stand-alone 18 stationary or mobile system and/or as an add-on to an existing recycle loop reactor 19 system 100 or to an existing cascade reactor system 320 to further increase the 20 percent conversion of methanol from these pre-existing reactors, or maintain the 21 overall conversion of the modified process while relaxing the effectiveness of the 22 recycle process through for example reduced recycle rate. Also, the system 52 23 can be used with systems 10 that produce other compounds formed via 24 equilibrium limited reactions, such as higher alcohols or dimethyl ether, among 25 others. 26 27 Various alternatives are contemplated as being within the scope of the 28 following claims particularly pointing and distinctly claiming the subject matter 29 regard as the invention. 27

Claims (12)

1. A method for the synthesis of a chemical compound formed by an equilibrium limited reaction, the method comprising the steps of: a) providing a reactor set including a first heat exchanger, and a reactor operably connected to the first heat exchanger and a second heat exchanger operably connected to the reactor, and a separator operably connected to the first and second heat exchangers, wherein at least one of the reactor or the separator is formed with a vessel adapted to withstand the operating temperatures and pressures of the reaction and including a number of zones formed within the vessel by dividing walls extending across the vessel to define the zones; each dividing wall adapted to withstand the temperature and pressure differentials across adjacent zones; b) directing a feed stream through the first heat exchanger formed as a plate-type or extended surface heat exchanger to preheat the feed stream; c) reacting the feed stream within the reactor; d) discharging a reaction product stream from the reactor through the first heat exchanger to preheat the incoming feed stream and to cool the reaction product stream; e) thermally contacting the reaction product stream with a cooling utility stream to form a condensed product within the reaction product stream; f) separating the condensed product from the reaction product; g) directing the reaction product stream through the second heat exchanger formed as a plate-type or extended surface heat exchanger to preheat the reaction product stream; and h) repeating steps b-g to achieve the desired product conversion.

2. The method of claim 1 where the feed stream is comprised of hydrogen and carbon oxide gases.

3. The method of claim 2 wherein the feed stream is a purge gas stream from a recycle stream of a chemical production reactor utilizing a steam reformer, and wherein the step 28 of directing the feed gas stream comprises directing the purge gas stream through the first heat exchanger.

4. A method for the synthesis of a chemical compound formed by an equilibrium limited reaction, the method comprising the steps of: a) providing a reactor set including a first heat exchanger and a reactor operably connected to the first heat exchanger and a second heat exchanger operably connected to the reactor; b) directing a feed stream through the first heat exchanger formed as a plate-type or extended surface heat exchanger to preheat the feed stream; c) reacting the feed stream within the reactor; d) discharging a reaction product stream from the reactor through the first heat exchanger to preheat the incoming feed stream and to cool the reaction product stream; e) thermally contacting the reaction product stream with a cooling utility stream to form a condensed product within the reaction product stream; f) separating the condensed product from the reaction product; g) directing the reaction product stream through the second heat exchanger formed as a plate-type or extended surface heat exchanger to preheat the reaction product stream; and h) repeating steps b-g to achieve the desired product conversion, wherein the first heat exchanger is formed with a first portion for thermally contacting the feed stream with the reaction product stream, and a second portion downstream from the first portion for thermally contacting the reaction product stream with the cooling utility, and wherein the step of thermally contacting the reaction product stream with the cooling utility takes place in the first heat exchanger.

5. The method of claim 4 wherein the first heat exchanger includes a third portion downstream from the second portion for separating the condensed product from the reaction product stream, and wherein the step of separating the condensed product from the reaction product stream takes place in the third portion of the first heat exchanger. 29

6. The method of claim 1 wherein the chemical compound is selected from the group consisting of: methanol, dimethyl ether and mixtures thereof

7. A method for condensing methanol out of a reaction product stream from a reactor of a methanol production system, the method comprising thermally contacting the reaction product stream from the reactor with a feed stream for the reactor in a plate type or extended surface heat exchanger to condense the methanol in the reaction product stream, wherein the reactor is formed with a vessel adapted to withstand the operating temperatures and pressures of the reaction and including a number of zones formed within the vessel by dividing walls extending across the vessel to define the zones; each dividing wall adapted to withstand the temperature and pressure differentials across adjacent zones.

8. A method for producing methanol comprising: a) providing a multiple reactor set including a reactor for converting a feed gas stream formed of a syn-gas created in a steam or autothermal reformer into a reaction product stream, a number of feed/effluent heat exchangers connected to the reactors and configured to preheat the feed gas stream, to cool the reaction product stream, and to condense the reaction product stream into a methanol stream and an excess feed gas stream, and a separator connected to the heat exchangers opposite the reactor for separating the methanol stream from the excess feed gas stream, wherein at least one of the reactor or the separator is formed with a vessel adapted to withstand the operating temperatures and pressures of the reaction and including a number of zones formed within the vessel by dividing walls extending across the vessel to define the zones; each dividing wall adapted to withstand the temperature and pressure differentials across adjacent zones; and b) condensing methanol from the reaction product stream at a rate of less than 2,500 tpd.

9. The method of claim 8 further comprising condensing methanol from the reaction product stream at a rate of less than 1,500 tpd. 30

10. The method of claim 8 further comprising condensing methanol from the reaction product stream at a rate of less than 1,000 tpd.

11. A method for producing a compound from a waste gas stream of a production system for the compound that employs a steam or autothermal reformer comprising the step of passing the waste stream through a multiple reactor set for converting a feed gas stream formed of the waste gas stream into a reaction product stream, a number of feed/effluent heat exchangers connected to the reactors and configured to preheat the feed gas stream, to cool the reaction product stream, and to condense the reaction product stream into a product compound stream and an excess feed gas stream, and a separator connected to the heat exchangers opposite the reactor set for separating the product compound stream from the excess feed gas stream, wherein at least one of the reactor or the separator is formed with a vessel adapted to withstand the operating temperatures and pressures of the reaction and including a number of zones formed within the vessel by dividing walls extending across the vessel to define the zones; each dividing wall adapted to withstand the temperature and pressure differentials across adjacent zones and wherein the multiple reactor set does not include a compressor.

12. The method of claim 11 wherein the wherein the feed stream is a purge gas stream from a chemical production reactor utilizing an autothermal reformer. 31