Abstract:
Methods and systems for heating a reactor feed in a multi reactor hydrocarbon dehydrogenation process. The methods and systems are advantageously employed for the production of styrene by the catalytic dehydrogenation of ethylbenzene. The catalytic dehydrogenation process employs heating steam operating at a steam to oil ratio of about 1.0 or less and relatively low steam superheater furnace temperature, such that all components exposed to steam in the process (outside of the fired heaters) can be constructed with standard metallurgy.

Description:
TECHNICAL FIELD 
     The various embodiments of the present invention relate to methods and systems having improved energy efficiencies in the production of styrene by dehydrogenation of ethylbenzene. The methods and systems reduce utility cost and provide savings in comparison with the current technology practiced in the industry. 
     BACKGROUND 
     It is well known in the art of styrene manufacture to react ethylbenzene (EB) in the presence of steam over a dehydrogenation catalyst, such as iron oxide under dehydrogenation reaction conditions, in order to strip hydrogen from the ethyl group on the benzene ring to form styrene. It is also well known that the dehydrogenation of ethylbenzene requires large amounts of energy, for example, in the form of steam. 
     Alternative methods for reducing energy consumption (i.e., steam) in processes for producing styrene via dehydrogenation of ethylbenzene have been previously described. 
     U.S. Pat. No. 4,628,136 to Sardina discloses a dehydrogenation process for producing styrene from ethylbenzene in the presence of steam by recovering heat of condensation normally lost during separation of the various components and using the heat to vaporize an aqueous feed mixture of ethylbenzene and water. Sardina teaches that this obviates the need to use steam to vaporize the liquid ethylbenzene feed. 
     U.S. Pat. No. 4,695,664 to Whittle discloses a means for recovering waste heat from a low temperature process stream with a vaporizable heat sink liquid and two immiscible liquids that form a low boiling azeotrope. The heat sink liquid is brought into indirect heat exchange with the low temperature process stream, whereby the heat sink liquid is able to recover heat from the process stream. 
     Various methods have been proposed that allow use of azeotropic heat recovery while operating at the minimum ratio of reaction steam to ethylbenzene, as determined by catalyst stability (i.e., resistance to coking). Such methods include use of direct heating as described in U.S. Pat. Nos. 8,193,404 and 8,084,660 to Welch et al., which discloses among other things methods for increasing the efficiency of a dehydrogenation unit by use of at least one direct heating unit. 
     Method of providing heat for chemical conversion and a process and system employing the method for the production of olefin to U.S. Pat. No. 8,163,971 to Wilcox et al. addresses the problem of supplying heat to the system at an overall steam/oil weight ratio of 1.0 or lower. Generally, these ratios would require steam temperature at the outlet of the steam superheater to be increased to 950° C., or even higher. At such high temperatures, the use of special and costly metallurgy is required. 
     U.S. Pat. No. 7,922,980 to Oleksy et al. discloses methods for recovering the heat of condensation from overhead vapor produced during ethylbenzene-to-styrene operations. In this regard, the &#39;980 patent uses the overhead of an EB/SM splitter column to vaporize an azeotropic mixture of ethylbenzene and water. 
     Other methods that could be employed to enable the use of azeotropic heat recovery while operating at the minimum ratio of reaction steam to ethylbenzene involve passing the reactor feed mixture through the convection section of a fired heater, as practiced by The Dow Chemical Company as described in U.S. Pat. No. 4,769,506 to Kosters. 
     Use of a split reheater arrangement as disclosed in published International Application No.: PCT/US2012/053100, Pub. No. W0/2014/035398, makes it possible to reduce the heating steam to ethylbenzene ratio required for interstage reheat to as low as 0.34 kg per kg of ethylbenzene. However, heating the primary reactor to a temperature required for efficient conversion of the ethylbenzene remains a separate problem. 
     Additionally, International Application No.: PCT/US2013/032244, Pub. No. W0/2014/142994, relates to efficiencies in the production of styrene through reduced quantities of steam used in the disclosed process. However, there still remains a need in the art for improvements that can provide even greater efficiencies through lower heating steam to ethylbenzene ratio, as presented herein. Without a means of supplying heat to the primary reactor feed prior to the addition of superheated steam, the temperature of the superheated steam added to the reactor feed upstream of the first reactor would exceed the mechanical temperature limits of the steam transfer line and the mixing device. To bring the temperature down, the amount of reaction steam has to be increased, which increases the overall energy demand of the process. 
     Thus, for economic reasons and process efficiencies, it is desirable to lower the reaction steam to hydrocarbon ratio of the process due to the costs incurred in generating and superheating steam. The inventive methods and systems disclosed herein provide for a reduction of reaction steam/EB ratio while practicing azeotropic heat recovery without resorting to the use of expensive alloys. 
     SUMMARY OF THE INVENTION 
     The various embodiments of the invention are directed to advantageously providing heat to processes to produce styrene monomer via ethylbenzene dehydrogenation, especially when the ratio of feed steam to ethylbenzene prior to the addition of superheated steam is greater than or equal to 0.4 by weight, as is the case when ethylbenzene is vaporized as an azeotropic mixture with water. 
     Consistent with the various embodiments of the present invention, a method of heating a reactor feed in a multi reactor hydrocarbon dehydrogenation process is disclosed, the method comprises the steps of: (i) heating a first reheating steam stream against flue gas from one or more fired heaters, after the reheating steam stream heats a first reactor product stream in one or more first reactor product stream indirect heating apparatuses, to provide a preheating steam stream; (ii) heating a first reactor feed stream comprising a hydrocarbon and optionally feed steam, with the preheating steam stream in a first reactor feed stream indirect heating apparatus prior to entering a first reactor, to provide a preheated hydrocarbon stream and a cooled preheating steam stream; (iii) heating the cooled preheating steam stream in one of the one or more fired heaters to provide a second reheating steam stream; (iv) directing the second reheating steam stream to one of the one or more first reactor product stream indirect heating apparatuses to provide heat to the first reactor product stream and a cooled reheating steam stream; (v) heating the cooled reheating steam in one of the one or more fired heater to provide a heating steam stream; and (vi) mixing the heating steam stream with the preheated hydrocarbon stream prior to entering the first reactor. 
     Consistent with the various embodiments of the present invention, a method of heating a reactor feed in a multi reactor hydrocarbon dehydrogenation process is disclosed, the method comprises the steps of (i) heating a first reactor feed stream comprising a hydrocarbon, and optionally feed steam, in an indirect heating apparatus with a preheating steam stream from one of one or more fired heater, to provide a preheated hydrocarbon stream and a cooled preheating steam stream; (ii) heating the cooled preheating steam stream against flue gas from the one or more fired heaters to provide a heating steam stream; and (iii) mixing the heating steam stream with the preheated hydrocarbon stream prior to entering a first reactor. 
     Consistent with the various embodiments of the present invention, a system for heating a reactor feed in a multi reactor hydrocarbon dehydrogenation process is disclosed, the system comprises a means of heating a steam stream used in a prior heat exchange step against flue gas from one or more fired heaters; and a means of indirectly transferring heat from said steam stream to a first reactor feed stream upstream of a first dehydrogenation reactor. 
     Consistent with the various embodiments of the present invention, a system for heating a reactor feed in a multi reactor hydrocarbon dehydrogenation process is disclosed, the system comprises a means of indirectly transferring heat from a steam stream to a first reactor feed stream in a heat transfer step upstream of a first dehydrogenation reactor to provide a preheated feed stream; and a means of heating the steam stream after the heat transfer step against flue gas from one or more fired heaters; wherein the steam stream, after reheating is mixed with preheated feed stream prior to entering the first dehydrogenation reactor. 
     The various improvements disclosed herein are substantial in tetras of their economic impact, i.e., through a reduction of energy consumption in the reaction section of a styrene manufacturing facility. Just as important, these improvements do not require fundamental changes to the process, for example, increased temperatures or pressures. 
    
    
     
       BRIEF DESCRIPTION OF THE DRAWING 
         FIG. 1  is a schematic flowsheet illustrating the reaction section of a two-reactor system for the production of styrene via dehydrogenation of ethylbenzene, with steam reheat. 
         FIG. 2  is a schematic flowsheet illustrating an embodiment of the invention having a two-reactor system and a series of two reheaters for producing styrene via dehydrogenation of ethylbenzene, with steam reheat. 
         FIG. 3  is a schematic flowsheet illustrating the reaction section of a two-reactor system for the production of styrene via dehydrogenation of ethylbenzene, with direct reheat. 
         FIG. 4  is a schematic flowsheet illustrating an embodiment of the invention having a two-reactor system for producing styrene via dehydrogenation of ethylbenzene, with direct reheat. 
         FIG. 5  is a schematic flowsheet illustrating an embodiment of the invention having a two-reactor system and a series of two reheaters for producing styrene via dehydrogenation of ethylbenzene, with steam reheat. 
         FIG. 6  is a schematic flowsheet illustrating an embodiment of the invention having a two-reactor system and two reheaters arranged in parallel for producing styrene via dehydrogenation of ethylbenzene, with steam reheat. 
         FIG. 7  is a schematic flowsheet illustrating an embodiment of the invention having a two-reactor system and two reheaters arranged in parallel for producing styrene via dehydrogenation of ethylbenzene, with steam reheat. 
     
    
    
     DETAILED DESCRIPTION OF THE INVENTION 
     Styrene is one of the most important monomers produced worldwide, and finds major use in the production of polystyrene, acrylonitrile—butadiene—styrene resins (ABS), and a variety of other polymers in the petrochemical industry. Styrene is produced commercially by catalytic dehydrogenation of ethylbenzene, and billions of pounds of styrene are produced each year. Therefore, the investment cost is very high, and even a small improvement in the plant operation can generate significant economic savings. Hence, optimal design and operation of the styrene producing equipment are required in styrene manufacturing processes. 
     Ethylbenzene dehydrogenation requires large amounts of energy in the form of steam. In particular, the dehydrogenation process requires large amounts of excess “reaction steam,” which is the total amount of steam needed to drive the endothermic reaction (i.e., the heat required to moderate the temperature drop as the reaction proceeds), reduce partial pressure of the reactants, and prevent catalyst coking. For the purposes of describing the methods and systems disclosed herein, the steam streams that ultimately constitute the steam that enters the dehydrogenation reactors (collectively referred as “reaction steam”) are defined as follows: (1) “feed steam” is steam that enters the process with the hydrocarbon feed, e.g., ethylbenzene feed, at the boundary limit of the system (i.e., the area demarcated by the dotted line in the Figures); (2) “preheating steam” is steam that is used to heat the hydrocarbon feed and feed steam mixture in an indirect heating apparatus, such as a shell-and-tube heat exchanger; (3) “reheating steam” is steam that is used to heat the effluent of a dehydrogenation reactor (before it enters another dehydrogenation reactor immediately downstream of it) in an indirect heating apparatus, and (4) “heating steam” is steam that is directly added to the ethylbenzene feed and feed stream mixture upstream of the first dehydrogenation reactor. As such, the reaction steam is used as a heat transfer medium to heat the reactor feeds in either heat exchangers and/or fired heaters, and feed steam which accompanies the ethylbenzene (EB) feed prevents coking in high temperature heat transfer equipment. 
     In the conventional ethylbenzene dehydrogenation process for producing styrene, a minimum of about 0.8 kg of heating steam per kg of ethylbenzene feedstock is required for two purposes: (1) reheating the feed steam and ethylbenzene feed between the primary and secondary reactors, which is needed because the dehydrogenation of ethylbenzene is a highly endothermic reaction; and (2) bringing the primary reactor feed steam and ethylbenzene feed mixture to the required reactor inlet temperature. 
     The ability to reduce the consumption of heating steam is particularly desirable for heat recovery schemes wherein an azeotropic mixture of ethylbenzene and water is boiled against the overhead of the EB/SM Splitter, or against the reactor effluent. The azeotropic mixture contains about 0.5 kg of feed steam per kg of ethylbenzene. An additional 0.10-0.15 kg of feed steam per kg of ethylbenzene can be generated by heat recovery from the reactor effluent, bringing the total feed steam to ethylbenzene ratio to 0.60-0.65 kg/kg. Since the minimum amount of reaction steam (sum of feed steam, preheating steam, reheating steam, and heating steam) is about 1 kg per kg of ethylbenzene, the amount of available heating steam is reduced by more than a factor of 2, i.e., from about 0.80 to about 0.35-0.40 kg per kg of ethylbenzene). 
     Even though prior art processes make it possible to reduce the heating steam to ethylbenzene ratio to the aforementioned low level for inter-stage reheating (for example, by utilizing the above-referenced split reheater arrangement), heating the feed steam and ethylbenzene feed to the primary dehydrogenation reactor to a temperature required for efficient conversion of the ethylbenzene remains a problem, the overcoming of which is the subject of the instant disclosure. Referring to  FIG. 1 , the prior art processes are limited because the feed steam and ethylbenzene feed stream  32  leave the feed/effluent exchanger  4  at a temperature that is significantly colder than the required reactor inlet temperature for primary dehydrogenation reactor  1 , typically by about 100° C. Thus, the minimum amount of heating steam stream  29  that is necessary to heat the feed steam and ethylbenzene feed to achieve the required temperature for the primary reactor feed stream  34  mixture is fixed by the mechanical limitations of the materials of construction of the steam transfer line for the heating steam stream  29  and primary reactor feed mixer  10 . In order to reduce the amount of heating steam for the styrene production process, additional means of heating the ethylbenzene feed is required. 
     Without a means of adding heat between where the feed steam and ethylbenzene feed leaves the feed/effluent exchanger  4  and the inlet of the primary dehydrogenation reactor  1 , heating steam temperatures in excess of 1000° C. are required when the heating steam to ethylbenzene ratio is reduced to less than 0.4 kg per kg of ethylbenzene while keeping the overall reaction steam to hydrocarbon ratio (ethylbenzene) no higher than 1.0. This is well beyond the limits of 800H/800HT metallurgy (899° C. as stated by ASME Code). In order to keep the steam temperature at a level where Alloy 800H/HT can be used, the heating steam flow must be roughly doubled, resulting in an overall reaction Steam-to-Oil ratio (S/O) in the reactors of about 1.25 kg per kg of ethylbenzene (EB) when azeotropic heat recovery is practiced. Since modern ethylbenzene dehydrogenation catalysts are capable of operating at S/O as low as 1.0, it is desirable to reduce the amount of heating steam. 
     The inventive flow schemes disclosed herein enable practice of azeotropic heat recovery in combination with low heating steam requirements, without the need for expensive and unproven materials of construction. 
     The Figures illustrate the differences between the current state of the art, i.e.,  FIG. 1  and  FIG. 3 , and the improved methods of this invention, i.e.,  FIG. 2  and  FIGS. 4-7 . Further, it is understood that certain equipment such as valves, piping, indicators and controls, and the like have been omitted from the Figures to facilitate the description thereof, and that the appropriate placement of such equipment is deemed to be within the scope of one skilled in the art. 
     Referring again to  FIG. 1 , the mixture of hydrocarbon and feed steam stream  31 , (hydrocarbon comprising e.g., ethylbenzene) which is either fully or partially vaporized, enters the shell side of the feed/effluent exchanger  4  where it is heated against the reactor effluent stream  38  from the secondary dehydrogenation reactor  3 . The feed steam and reactor feed stream  32  is then directed to the primary reactor feed mixer  10  where it is mixed with the heating steam stream  29  coming from the tertiary steam superheater  7  (steam superheaters as known in the art). The temperature of the superheated steam is adjusted to achieve the desired/required temperature for total primary reactor feed stream  34  mixture, which is transferred to the primary dehydrogenation reactor  1 . 
     Due to the endothermic nature of the dehydrogenation reaction, multiple reactors are required to effect the significant ethylbenzene conversion needed to make the process economic. The inter-stage reheating is typically accomplished in a shell and tube heat exchanger utilizing steam as a heating medium. This type of reheater (i.e., an indirect heating apparatus) is commonly referred to as steam reheater. The reheating can also be accomplished using so called direct heating methods, which include the use of Flameless Distributed Combustion (described in U.S. Pat. No. 8,084,660), or selective oxidation of hydrogen (UOP-Lummus SMART process). 
       FIGS. 1 and 2  represent a high temperature chemical conversion processes, which utilize two catalytic reactors as conversion apparatuses, i.e., primary (first) dehydrogenation reactor  1  and secondary (second) dehydrogenation reactor  3 , respectively, using steam reheat, wherein the reheat duty is split between a series of two reheaters, i.e., primary reheater  2 A and secondary reheater  2 B, respectively, to provide product stream  39 . The reactor effluent stream  35  from primary dehydrogenation reactor  1  is first heated in a series primary reheater  2 A against reheating steam stream  23  coming from the primary steam superheater  5  (a fired heater), to produce the reactor effluent stream  36  which is then heated further on the tube side of series secondary reheater  2 B against reheating steam stream  27  coming from the secondary steam superheater  6 . The temperature of the steam delivered to the primary and secondary reheaters  2 A and  2 B is adjusted to achieve the desired temperature of process feed stream  37  to the secondary dehydrogenation reactor  3 . 
     The feed stream  31  of  FIGS. 1 through 4  provides a process stream for the conversion of a hydrocarbon component in a high temperature chemical conversion process. The process stream, for example, comprises steam and a hydrocarbon component, such as, ethylbenzene that is converted to styrene in the high temperature chemical conversion process. 
     Further in  FIGS. 1 and 2 , low temperature steam stream  21  (which ultimately becomes heating steam in all its various applications) is first preheated against the combined flue gas stream  13 , i.e., flue gas, (and  13 A in  FIG. 2 )) from the primary, secondary, and tertiary steam superheaters  5 ,  6  and  7  in primary convective heating coil  8 , located in the convection section shared by the superheaters  5 ,  6  and  7 . The steam then passes as stream  22  through the radiant coil of the primary steam superheater  5 , and from there it is delivered as reheating steam stream  23  to the primary reheater  2 A. 
     In  FIG. 1 , the cooled primary reheating steam stream  24  (i.e., after it leaves primary reheater  2 A) is heated in the radiant coil of the secondary steam superheater  6  before being sent as secondary reheating steam stream  27  to the secondary reheater  2 B. Finally, after leaving reheater  2 B the secondary reheating steam stream  28  passes through the radiant coil of the tertiary steam superheater  7  before being added as heating steam stream  29  to the first reactor feed stream  32  in primary reactor feed mixer  10  to bring the resulting final mixture of primary reactor feed stream  34  to the required temperature before entering primary reactor  1 . 
     In  FIGS. 1 and 2 , heat to the superheaters  5 ,  6  and  7  is supplied by combustion of fuel  11 A in air  12 A, which provides for combined flue gas  13  (and combined flue gas  13 A in  FIG. 2 ). The flue gas  15  leaving primary convective heating coil  8  is directed to an economizing coil  9  prior to being vented to atmosphere through stack  16 . 
     Some of the difference between the prior art process of  FIG. 1 , as described above and the processes of the various embodiments of the instant invention, as presented in  FIG. 2 , comprises the addition of a secondary convection heating coil  11  and a reactor feed preheater  12  apparatus (i.e., an indirect heating apparatus). Cooled reheating steam stream  24  exiting the primary reheater  2 A is heated against the combined superheater flue gas streams  13  in secondary convection heating coil  11  and then directed to the shell side of the process feed preheater  12  as preheating steam stream  25  where it exchanges heat with the reactor feed stream  32 . Cooled preheating steam stream  26  leaves process feed preheater  12  and is directed to the radiant coil of the secondary steam superheater  6  before being sent as secondary reheating steam stream  27  to secondary reheater  2 B. 
     In  FIG. 2 , since the temperature of the preheated feed stream  33  (i.e., heated hydrocarbon stream) entering the primary reactor feed mixer  10  is higher than that of corresponding reactor feed stream  32  in the prior art process of  FIG. 1 , the temperature of the superheated heating steam stream  29  necessary to achieve the same temperature in the primary reactor feed stream  34  (i.e., a mixed reactor feed stream) mixture at the inlet to the primary dehydrogenation reactor  1  is reduced compared with the process of prior art  FIG. 1 . Alternatively, keeping the tertiary steam superheater  7  outlet temperature the same as in the current state of the art process, the amount of heating steam required is reduced. This reduction in heating steam is illustrated by Example 1 in Table 1. 
     As presented in Table 1, the process data of Comparative Example 1A, represent the schematic flow sheet of  FIG. 1 , i.e., the prior art process. Comparative Example 1A has an overall steam-to-ethylbenzene ratio is 1.00 kg/kg, and 62% of the total steam used in the reactors is generated in the process itself by a combination of azeotropic heat recovery and steam generation utilizing the heat of the reactor effluent. Thus, only 0.38 kg of heating steam per kg of ethylbenzene feed is available for interstage reheat and for heating the primary reactor feed. In Comparative Example 1A, the primary dehydrogenation reactor  1  inlet temperature necessary to convert 64% of ethylbenzene in the two reactor system is 650° C. In Comparative Example 1A, the mixture of ethylbenzene and feed steam leaves the feed/effluent exchanger  4  at 550° C. By heat balance, the heating steam has to be heated in the tertiary steam superheater  7  to a temperature of 1102° C. in order for the final primary reactor feed stream  34  mixture to reach the required 650° C. primary dehydrogenation reactor  1  inlet temperature. However, the 1102° C. temperature is well above the temperature limit of the Incoloy 800H/800 HT metallurgy (899° C.), used for the construction of the heating steam stream  29  transfer line and the primary reactor feed mixer  10 . 
     In order to reduce the temperature of superheated heating steam stream  29  to 899° C., the amount of heating steam has to be increased from 0.38 kg to 0.63 kg per kg of ethylbenzene, which increases the operating cost, both due to the cost of the additional steam and the cost of additional fuel necessary to heat the steam. The data presented in Table 1 for the prior art process of Comparative Example 1B utilize the increased quantity of heating steam, i.e., 0.63 kg per kg of ethylbenzene, which is necessarily required to reduce the temperature of superheated heating steam stream  29  to 899° C. 
     
       
         
               
               
               
               
             
               
               
               
               
             
           
               
                 TABLE 1 
               
               
                   
               
               
                   
                   
                 Comparative 
                 Comparative 
               
               
                   
                 Example 1 
                 Example 1A 
                 Example 1B 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 Overall Steam/Hydrocarbon Ratio (kg/kg) 
                 1.00 
                 1.00 
                 1.25 
               
               
                 Heating Steam/Hydrocarbon Ratio (kg/kg) 
                 0.38 
                 0.38 
                 0.63 
               
               
                 Hydrocarbon Flow (kg/hr) 
                 107345 
                 107345 
                 107345 
               
               
                 Stream 21 Temperature (° C.) 
                 155 
                 155 
                 155 
               
               
                 Stream 22 Temperature (° C.) 
                 448 
                 550 
                 441 
               
               
                 Stream 23 Temperature (° C.) 
                 850 
                 850 
                 810 
               
               
                 Stream 24 Temperature (° C.) 
                 596 
                 596 
                 619 
               
               
                 Stream 25 Temperature (° C.) 
                 789 
                   
                   
               
               
                 Stream 26 Temperature (° C.) 
                 565 
                   
                   
               
               
                 Stream 27 Temperature (° C.) 
                 899 
                 899 
                 772 
               
               
                 Stream 28 Temperature (° C.) 
                 643 
                 643 
                 648 
               
               
                 Stream 29 Temperature (° C.) 
                 899 
                 1102 
                 899 
               
               
                 Stream 32 Temperature (° C.) 
                 550 
                 550 
                 550 
               
               
                 Stream 33 Temperature (° C.) 
                 597 
                   
                   
               
               
                 Stream 34 Temperature (° C.) 
                 650 
                 650 
                 641 
               
               
                 Stream 35 Temperature (° C.) 
                 560 
                 560 
                 559 
               
               
                 Stream 36 Temperature (° C.) 
                 605 
                 605 
                 609 
               
               
                 Stream 37 Temperature (° C.) 
                 650 
                 650 
                 641 
               
               
                 Stream 38 Temperature (° C.) 
                 588 
                 587 
                 585 
               
               
                 Primary Radiant Coil (5) Duty (10 6  kcal/hr) 
                 8.78 
                 6.65 
                 13.16 
               
               
                 Secondary Radiant Coil (6) Duty (10 6  kcal/hr) 
                 6.83 
                 6.83 
                 5.60 
               
               
                 Tertiary Radiant Coil (7) Duty (10 6  kcal/hr) 
                 5.81 
                 10.73 
                 9.37 
               
               
                 Primary Convective Coil (8) Duty (10 6  kcal/hr) 
                 5.97 
                 8.11 
                 9.59 
               
               
                 Secondary Convective Coil (11) Duty (10 6  kcal/hr) 
                 4.28 
                   
                   
               
               
                 Economizing Coil (9) Duty (10 6  kcal/hr) 
                 4.52 
                 12.59 
                 7.19 
               
               
                 Superheater Fuel Consumption (10 6  kcal/hr) 
                 39.34 
                 48.81 
                 48.80 
               
               
                 External Steam Consumption (10 6  kcal/hr) 
                 22.53 
                 22.53 
                 37.03 
               
               
                 Total Energy Consumption (10 6  kcal/hr) 
                 61.87 
                 71.33 
                 85.83 
               
               
                 Net Energy Consumption (after Coil 11 credit) (10 6   
                 57.35 
                 58.74 
                 78.65 
               
               
                 Annual Energy Cost at $5-15 per 10 6  kcal (million USD) 
                 9.2-27.5 
                 9.4-28.2 
                 12.6-37.8 
               
               
                 Annual Savings (million USD) 
                 3.4-10.2 
                 N/A 
                 Base 
               
               
                   
               
             
          
         
       
     
     On the other hand, the process data of Example 1 presented in Table 1 represent the inventive schematic flow sheet of  FIG. 2 . The process of  FIG. 2  is different from prior art  FIG. 1  in that the claimed flow process inventively utilizes a process feed preheater  12 , and a secondary convection heating coil  11 . The process feed preheater  12  preheats reactor feed stream  32  with preheating steam stream  25 . Because the reactor feed stream  33  entering the feed mixer  10  is at a higher temperature than the rector feed steam  32  in the existing state of the art process, the heating temperature required to bring the first reactor feed to the desired temperature at the inlet to the first reactor is reduced in the methods and systems disclosed herein. Preheating steam stream  25  is ultimately provided by utilizing primary reheating steam stream  23  from primary steam superheater  5  in a heat transfer step to reheat the reactor effluent stream  35  (i.e., a high temperature chemical conversion apparatus effluent), in primary reheater  2 A to provide cooled reheater steam stream  24 , which is then forwarded to secondary convection heating coil  11 . Secondary convection heating coil  11  advantageously utilizes the heat from combined flue gases  13  of primary, secondary, and tertiary steam superheaters  5 ,  6 , and  7 . Once preheating steam stream  25  exchanges heat with the reactor feed stream  32  in process feed preheater  12 , it becomes cooled preheating steam stream  26 , which is forwarded to secondary steam superheater  6  to provide secondary reheating steam stream  27 . 
     In particular, the present inventive method and system, make it possible to keep the overall steam-to-ethylbenzene ratio at 1.00 kg/kg without violating the temperature limit of the Incoloy 80011/800 HT material because the temperature of the reactor feed entering the primary reactor feed mixer  10  is raised from 550° C. to 597° C. in the process feed preheater  12 . 
     From the data presented in Table 1, in a facility producing 500 thousand metric tons of styrene annually, the net energy savings compared with the current state of the art process (after accounting for the energy recovered in the economizer coil  9 ) are 21.3·10 6  kcal/hour. This translates to annual savings of between 3.4 and 10.2 million USD (using a range of fuel prices between 20 and 60 USD per million kcal). 
     As mentioned earlier, the methods of this invention can also be applied to a system using direct heating for interstage reheat. In such a system, interstage reheating is accomplished by burning fuel in oxygen or air directly inside the reheater, as described in U.S. Pat. No. 8,084,660 to Welch et al., by selectively oxidizing hydrogen which comprises part of the primary reactor effluent, as practiced by the UOP-Lummus SMART process. 
       FIG. 3  depicts a flow scheme of the prior art in which direct heating is used, as described above. In  FIG. 3  the low temperature steam stream  21  (which ultimately becomes heating steam) is first heated against the superheater flue gas stream  13  and is then heated to the final temperature in the radiant coil of the steam superheater  5 , before being added as heating steam stream  29  to the reactor feed stream  32  in the primary reactor feed mixer  10 . 
       FIG. 4  depicts a flow scheme embodiment of the invention as applied to a system using direct heating for interstage reheat. The key difference between the prior art process of  FIG. 3 , as described above, and the process of the various embodiments of the invention is the addition of a process feed preheater  12  and a secondary convection coil  11 . The process feed preheater  12  takes the preheating steam stream  23  and uses it to preheat the reactor feed stream  32  to an intermediate temperature. The cooled preheating steam stream  26  from the process feed preheater  12  is directed to the secondary convection coil  11 . The preheated feed stream  33  leaving the process feed preheater  12  is then brought to the primary reactor feed stream  34  mixture temperature by mixing it in the primary reactor feed mixer  10  with the heating steam stream  29  heated in the secondary convection coil  11 . The feed stream is then reacted in primary reactor  1 , before being passed as effluent stream  35  to reheater  2  and on to secondary dehydrogenation reactor  3  via effluent stream  37 , to provide the final reactor product stream  38 , which after exchanging heat with the reactor feed stream  31  becomes stream  39 . 
       FIG. 5  presents an embodiment of the invention having a two-reactor system and a series of two reheaters for producing styrene via dehydrogenation of ethylbenzene.  FIG. 5  presents the option of reheating steam stream  23  being directed to the series of two reheaters and specifically reheater  2 B instead of reheater  2 A. Reheating steam stream  27  coming from the fired heater  6  (i.e., secondary steam superheater) is directed to reheater  2 A instead of reheater  2 B. 
       FIG. 6  presents an embodiment of the invention having a two-reactor system and two reheaters arranged in parallel.  FIG. 6  presents the option of reheating steam stream  23  directed to reheater  2 A and reheating steam stream  27  is directed to reheater  2 B.  FIG. 7  presents an embodiment of the invention having a two-reactor system and two reheaters arranged in parallel with the option of reheating steam stream  23  directed to reheater  2 B and reheating steam stream  27  being directed to reheater  2 A. 
       FIGS. 2 and 4  differ from one another as far as the source of steam used in the process feed preheater  12  and the destination of the steam leaving the secondary convection coil  11 . In the case of a steam reheat based system of  FIG. 2 , the preheating steam used in the process feed preheater  12  comes from the secondary convection coil  11 , while in the case of a direct heating based system of  FIG. 4  the preheating steam comes from the primary steam superheater  5 , and heating steam is heated in the secondary convection heating coil  11  before being directed to the primary reactor feed mixer  10 . 
     The process data presented in Table 2 compare Comparative Examples 2A and 2B to an embodiment of the invention of Example 2. The Examples of Table 2 represent a system wherein a high amount of steam is generated in the process itself by a combination of azeotropic heat recovery and steam generation utilizing the heat of the reactor effluent. As such, the process parameters are identical to those in the Examples presented in Table 1, i.e., the overall steam-to-ethylbenzene ratio is 1.00 kg/kg, with 62% of the total steam used in the reactors being generated in the process itself by a combination of azeotropic heat recovery and steam generation using the heat of the reactor effluent, and the reactor feed being heated to 550° C. in the feed/effluent exchanger  4 . 
     
       
         
               
               
               
               
             
               
               
               
               
             
           
               
                 TABLE 2 
               
               
                   
               
               
                   
                   
                 Comparative 
                 Comparative 
               
               
                   
                 Example 2 
                 Example 2A 
                 Example 2B 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 Overall Steam/Hydrocarbon Ratio (kg/kg) 
                 1.00 
                 1.00 
                 1.25 
               
               
                 Heating Steam/Hydrocarbon Ratio (kg/kg) 
                 0.38 
                 0.38 
                 0.63 
               
               
                 Hydrocarbon Flow (kg/hr) 
                 107345 
                 107345 
                 107345 
               
               
                 Stream 21 Temperature (° C.) 
                 155 
                 155 
                 155 
               
               
                 Stream 22 Temperature (° C.) 
                 350 
                 550 
                 400 
               
               
                 Stream 23 Temperature (° C.) 
                 899 
                   
                   
               
               
                 Stream 24 Temperature (° C.) 
                   
                   
                   
               
               
                 Stream 25 Temperature (° C.) 
                   
                   
                   
               
               
                 Stream 26 Temperature (° C.) 
                 596 
                   
                   
               
               
                 Stream 27 Temperature (° C.) 
                   
                   
                   
               
               
                 Stream 28 Temperature (° C.) 
                   
                   
                   
               
               
                 Stream 29 Temperature (° C.) 
                 820 
                 1104 
                 899 
               
               
                 Stream 32 Temperature (° C.) 
                 550 
                 550 
                 551 
               
               
                 Stream 33 Temperature (° C.) 
                 615 
                   
                   
               
               
                 Stream 34 Temperature (° C.) 
                 650 
                 650 
                 641 
               
               
                 Stream 35 Temperature (° C.) 
                 561 
                 561 
                 560 
               
               
                 Stream 36 Temperature (° C.) 
                   
                   
                   
               
               
                 Stream 37 Temperature (° C.) 
                 650 
                 650 
                 641 
               
               
                 Stream 38 Temperature (° C.) 
                 588 
                 588 
                 585 
               
               
                 Primary Radiant Coil (5) Duty (10 6  kcal/hr) 
                 10.08 
                 12.78 
                 17.94 
               
               
                 Secondary Radiant Coil (6) Duty (10 6  kcal/hr) 
                   
                   
                   
               
               
                 Tertiary Radiant Coil (7) Duty (10 6  kcal/hr) 
                   
                   
                   
               
               
                 Primary Convective Coil (8) Duty (10 6  kcal/hr) 
                 3.99 
                 8.10 
                 8.21 
               
               
                 Secondary Convective Coil (11) Duty (10 6  kcal/hr) 
                 4.97 
                   
                   
               
               
                 Economizing Coil (9) Duty (10 6  kcal/hr) 
                 1.54 
                 5.22 
                 4.47 
               
               
                 Superheater Fuel Consumption (10 6  kcal/hr) 
                 22.37 
                 28.37 
                 33.29 
               
               
                 External Steam Consumption (10 6  kcal/hr) 
                 22.53 
                 22.53 
                 37.03 
               
               
                 Total Energy Consumption (10 6  kcal/hr) 
                 44.90 
                 50.89 
                 70.32 
               
               
                 Net Energy Consumption (after Coil 11 credit) (10 6   
                 43.35 
                 45.67 
                 65.84 
               
               
                 Annual Energy Cost at $5-15 per 10 6  kcal (million USD) 
                 6.9-20.8 
                 7.3-21.9 
                 10.5-31.6 
               
               
                 Annual Savings (million USD) 
                 3.6-10.8 
                 N/A 
                 Base 
               
               
                   
               
             
          
         
       
     
     As expected, the Comparative Examples 2A and 2B, which represent the current state of the art methods, suffer from the same limitations as in Comparative Examples 1A and 1B, respectively. Specifically, the amount of heating steam available is insufficient to keep the temperature of the heating steam stream  29  at 899° C. or less. However, the improved method of Example 2 as exemplified in  FIG. 4 , preheats the reactor feed to 615° C. with 899° C. preheating steam originating in the primary steam superheater  5 , and the former is then brought to the required 650° C. reactor inlet temperature by addition of heating steam heated in the secondary convection coil  11 . 
     The methods of the claimed invention are equally well applicable to a system consisting of a single or multiple dehydrogenation reactors. The reactors can be adiabatic or isothermal, radial flow or axial flow type, or any combination of these characteristics. Furthermore, the methods of this invention are equally well applicable to styrene reaction systems featuring steam reheat and those featuring direct heating for interstage reheat, such as UOP-Lummus SMART process and the process disclosed in U.S. Pat. Nos. 8,193,404 and 8,084,660 to Welch et al. Also, the methods of this invention can be used with any catalyst suitable for dehydrogenation of ethylbenzene. 
     The terms “invention,” “the invention,” “this invention,” and “the present invention” used in this patent are intended to refer broadly to all of the subject matter of this patent and the patent claims below. Statements containing these terms should not be understood to limit the subject matter described herein or to limit the meaning or scope of the patent claims below. 
     Although the embodiments of the present invention have been described in considerable detail with regard to certain versions thereof, other versions are possible, and alterations, permutations, and equivalents of the version shown will become apparent to those skilled in the art upon a reading of the specification and study of the drawings. Also, the various features of the versions herein can be combined in various ways to provide additional versions of the present invention. Furthermore, certain terminology has been used for the purposes of descriptive clarity, and not to limit the present invention. Therefore, any appended claims should not be limited to the description of the preferred versions contained herein and should include all such alterations, permutations, and equivalents as fall within the true spirit and scope of the present invention.