BACKGROUND OF THE INVENTION

Field of the invention The present invention relates to a distillation-reaction system wherein a portion of the heat of reaction is recovered and used to operate the distillation. Related information In combined reaction/distillation systems, such as the catalytic distillation system, where the catalyst is present in a form suitable to serve as a distillation structure; the heat of reaction generated in the reaction zone is utilized to reduce the energy requirement of the distillation process, i.e., reboiler duty. This is indeed the case when the key separations occur in staging within and above the reaction zone. On the other hand if key separations occur in staging below the reaction zone, heat of reaction does not reduce the heat duty to the column but simply increases condenser cooling duty. An advantage of this invention is a flexible and efficient method that makes use of the heat of reaction to reduce reboiler duty in reaction-distillation systems in which the key separation occurs prior to reaction. SUMMARY OF THE INVENTION Briefly the present invention is a method to recover the heat of reaction of a reaction component from a petroleum stream to assist in fractionating the petroleum stream to remove and recover the reaction component. The present invention includes the process for recovering the heat of reaction of a reaction component from a multi component petroleum stream to assist in fractional distillation of the multi component petroleum stream from which the reaction component is derived comprising: fractionating said multi component petroleum to recover at least two fractions, a first fraction containing said reaction component having a first boiling range and comprising less than the entire multi component petroleum stream, and a second fraction having a second boiling point which is lower than the first boiling range; removing said first fraction from said fractional distillation; reacting said reaction component under conditions to exothermically react said reaction component and produce a reaction stream having a third temperature higher than said second temperature range; removing said second fraction from said fractional distillation; heating said second fraction by indirect heat exchange with said reaction stream, to heat said second fraction and returning the second fraction to fractional distillation. Preferably there are at least two fractions having a boiling range less than said first fraction, which are removed from the fractional distillation and heated by indirect contact with said reaction stream. Preferably the highest boiling range fraction other than said first fraction removed from the fractional distillation is first contacted with said reaction stream and each fraction other than said first fraction is contacted indirectly with said reaction stream in order of their descending boiling range and thereafter returned to the fractional distillation. DETAILED DESCRIPTION The invention provides an energy efficient process scheme for a highly exothermic distillation reaction system wherein the reactor is external to the distillation column and wherein feed to the reactor comprises a mixture of at least one liquid product stream from the distillation column. The reactor is operated under adiabatic and boiling point conditions and at a pressure that results in vaporizing a portion of the liquid flow through the reactor due to the heat of reaction. Under these conditions reaction temperature is controlled by reactor pressure. The invention contemplates that the pressure (and hence the temperature) will be maintained at a sufficiently high level such that the reactor effluent can be efficiently used to provide reboil heat for the distillation column. In one embodiment of the invention the reactor operates in plug down flow mode, reactor effluent is routed to a column side reboiler, and heat utilization is accomplished by maintaining the reactor pressure at a higher level than the distillation column such that the reactor effluent dew point temperature range is higher than the distillation column reboiler temperature. Following are examples of industrial processes that would benefit from one or more aspects of this invention, wherein the feed to reaction is derived from a prior fractional distillation: reduction of benzene content in reformate streams slated for motor gasoline use by converting the benzene to cyclohexane via hydrogenation; cyclohexane via hydrogenation of benzene; aniline via hydrogenation of nitrobenzene; aromatics alkylation, e.g., ethylbenzene, cumene, butyl benzene; oxidation systems, e.g., vinylacetate via ethylene/02/acetic acid; and hydroformylation systems, e.g., Fischer-Tropsch products; methanol via CO/H2. The figure is a generalized implementation of a process for converting a compound A contained in a feed stream with lighter and heavier components with a compound B to form a product C in an exothermic reaction in reactor 2. The overall process scheme consists of distillation column 1 ; reaction zone 2; liquid-vapor separators 3, 4; reboilers 6, 7, 8; heater 9; condensers 10, 11 ; feed/effluent exchanger 12. Heat sources for reboiler 6 may be steam, hot oil, or process heaters. Heat sink for condensers 10 and 11 may be cooling water or air. Heat sources for the remaining heat exchangers are obtained from the heat of reaction as described below. Feed stream 101 a to the distillation column is a multicomponent petroleum which is comprised of heavy-end components, intermediate boiling range components including compound A, light boiling range components, and light end components. The column 1 , configured with mass transfer stages above and below the feed point, is designed to split the feed into its light boiling range components in stream 102, intermediate range components including essentially all of compound A in stream 104 which is withdrawn from the column several stages above the feed stage, and heavy end components in stream 103. This is accomplished by providing sufficient stages above and below the feed stage and heat input to reboilers 6, 7, and 8. The heavy-end product stream 107 is obtained after cooling stream 103 by heat exchange with feed stream 101 a. Reactor 2 is an adiabatic boiling point reactor containing catalyst that promotes the desired reaction. Feed to the reactor includes stream 104 (via heater 9), stream 108 containing compound B, and recycle stream 114 generated as noted below. Stream 104 enriched in compound A and essentially free of heavy-end components is withdrawn from several stages above the feed stage. Stream 108 can be either vapor or liquid. The reactor 2 is configured as a typical fixed bed reactor operating up-flow or down-flow or in a preferred embodiment as a down-flow reactor containing catalyst supported within a mass transfer structure. Heat is generated in the reactor due to the heat of reaction. Since the reactor is operating under adiabatic and boiling point conditions, a portion of the liquid phase flowing through the reactor vaporizes in an amount corresponding to the heat of reaction and latent heat of vaporization. In one aspect of the invention, sufficient liquid flow to the reactor is provided so that the reactor effluent, stream 110, comprises both liquid and vapor phases. This is accomplished by adjusting the flow of recycle stream 114 and the pressure of reactor 2. In one aspect of the invention, stream 110 is routed to the hot side inlet of reboiler 7 where it is heat exchanged against side draw stream 106 fed to the cold inlet side of reboiler 7. Hot and cold side exit streams are streams 111 and 105 respectively. In another aspect of the invention, reactor 2 Operates at a higher pressure than column 1 such that the temperature of stream 111 is higher than the bubble point of stream 106. Underthis condition, a portion of the sensible and latent heat in stream 110 go to boil stream 106 producing vapor stream 105 which is fed back into column 1. In another aspect of the invention the difference in operating pressure between reactor 2 and column 1 is adjusted such that the resulting temperatures of the feed and effluent streams provide at least a two degree Fahrenheit temperature approach at either the inlet or outlet sides of reboiler 7. In another aspect of the invention stream 111 is optionally routed to the hot side inlet of reboiler 8 which provides a portion of the reboiler duty to the lights intermediate component section of column 1. Cold side inlet is stream 115 drawn from the feed stage. Cold side exit is vapor stream 116 returned to column 1 above the draw stage. Hot side exit stream 112 is fed to vapor-liquid separation vessel 3 producing vapor stream 117 and liquid stream 113. Stream 113 is split into recycle stream 114 and product stream 121 containing intermediate boiling range components including any unconverted compound A, reaction product compound C, and any condensable components in stream 102 including unconverted reactant B. Stream 114/stream 121 split ratio is an independent process variable which in conjunction with operating pressure of the reactor controls the temperature profile across the reactor as well as the vapor-liquid flow distribution. In a preferred embodiment of the invention the combination of reactor pressure and recycle stream flow is adjusted such that the reactor is operating in pulse flow mode (mass velocities of liquid and vapor generally >3000 Ib/h/ft2) and the corresponding temperature profile results in practical space yield, conversion and selectivity for reaction. In still another aspect of the invention, remaining sensible and latent heat in stream 117 is optionally used to preheat stream 104 to stream 104a feed to reactor 2 by routing streams 104 and 117 to the cold and hot inlet sides respectively of heater 9. Stream 118 containing light non-condensing components that were either produced in the reactor or were contained in feed stream 108 is vapor/liquid separated in vessel 4. The vapor stream is further cooled in condenser 11 producing condensate stream 119 which is returned to vessel 4 and vent stream 120 comprising light end components that entered the system in stream 108 and any light end by-products produced in the reactor. The combination of streams 119 and 121 is the intermediate boiling range product stream 122 which contains product C. Another aspect of the invention (configuration not shown) addresses the situation where the volatility of product component C is in the range of the heavy boiling range components. In this case stream 122 is fed back to column 1 to separate product C as a liquid bottoms product or as a vapor product drawn from a stage near the bottom of the column while intermediate boiling range components are recovered from the column together with light-end components in stream 102. EXAMPLE 1 The following example demonstrates use of the invention for reducing the benzene content in a gasoline mixture containing C5-C10 paraffin and aromatic components. Equipment and stream names are as given in the figure. Compositions and stream flows are in Table 1. Main reaction in reactor 2 is the following reaction catalyzed by a supported Ni catalyst: Benzene + 3 Hydrogen → cyclohexane Reactor operating conditions are summarized in Table 2. TABLE 2 Pressure in 250 psi Pressure out 245 psi Temperature in 2230F Temperature out 358°F LHSV(based on total liquid feed) 6.9 ft3 liquid feed/hr/ ft3 reactor volume Benzene/H2 mol ratio in feed 3.51 Mass flow liquid in 10,850 Ib/hr/ft2 Mass flow liquid out 5,683 Ib/hr/ft2 Mass flow vapor in 1 ,832 Ib/hr/ft2 Mass flow vapor out 7,000 Ib/hr/ft2 Benzene conversion across reactor 99 % Reactor recvcle mass ratio. 2.36 stream 9/stream 15 Column 1 design parameters are summarized in Table 3. Separation requirements are less than 1 wt% toluene in stream 104 (to minimize toluene loss in reactor 2 by hydrogenation to methylcyclohexane) and benzene levels in streams102 and 103 of less than 0.5 wt%. TABLE 3 Number of stages 62 Top pressure 29 psi Reflux mass ratio, reflux rate/distillate rate 7.83 Feed stream locations Stream 151 Stage 37 Stream 155 Stage 57 Stream 116 Stage 26 Draw stream locations Stream 153 Stage 62 Stream 115 Stage 26 Stream 156 Stage 57 Stream 152 Stage 1 Stream 154 Stage 26 Reboiler duties Reboiler 6 (external heat source) 23 MM btu/hr Reboiler 7 9.8 MM btu/hr Reboiler β 8.0 MM btu/hr The reboiler data in Table 3 demonstrate that practice of the invention reduces external heat load to the distillation column by 43.6% corresponding to use of 82% of the heat of reaction generated in the reactor. Table 1 - part 1 STREAM 101 101a 102 103 104 104a 105 106 107 108 109 1 10 111 112 Temperature F 257 176 123 312 201 254 356 290 194 86 223 359 320 268 Pressure psi 145 145 29 31 30 250 40 31 31 250 250 245 245 245 Mass Flow Ib/hr H2 0 0 0 0 0 0 0 0 0 1649 1660 251 251 251 Methane 0 0 0 0 0 0 0 0 0 1680 1952 1952 1952 1952 Ethane 0 0 0 0 0 0 0 0 0 1889 2626 2626 2626 2626 Propane 0 0 0 0 0 0 0 0 0 1709 2878 2878 2878 2878 I-Butane 799 799 777 0 23 23 0 0 0 0 . 44 44 44 44 N-Butane 1865 1865 1794 0 71 71 0 0 0 791 1784 1784 1784 1784 I-C5 11721 11721 10730 0 991 991 0 0 0 378 3243 3243 3243 3243 N-C5 8258 8258 7398 0 860 860 0 0 0 302 2843 2843 2843 2843 2,2 Dimethylbutane 1865 1865 1476 0 389 389 0 0 0 0 1015 1015 1015 1015 2,3-Dimethylbutane 2131 2131 1328 0 803 803 0 0 0 0 2159 2159 2159 2159 OO Cycpentane 0 0 0 0 0 0 0 0 0 0 0 0 0 0 Cycpentene 266 266 226 0 41 41 0 0 0 0 103 103 103 103 1-Pentene 266 266 242 0 25 25 0 0 0 0 59 59 59 59 2-Methylpentane 7459 7459 3890 0 3569 3569 0 0 0 0 9632 9632 9632 9632 3-Methylpentane 6127 6127 1552 0 4575 4575 0 0 0 0 12489 12489 12489 12489 N-Hexane 7459 7459 146 0 7313 7313 1 1 0 0 20312 20312 20312 20312 1-Hexene 266 266 73 0 193 193 0 0 0 0 526 526 526 526 Cychexene 533 533 0 81 452 452 102 102 81 0 1302 1302 1302 1302 Cychexane 0 0 0 0 0 0 0 0 0 0 36946 56567 56567 56567 2,2-Dimethylpentane 799 799 0 0 796 796 7 7 3 0 2281 2281 2281 2281 Methylcyclopentane 1066 1066 5 1 1060 1060 3 3 1 0 2976 2976 2976 2976 2-4-Dimethylpentane 799 799 0 0 796 796 9 9 4 0 2283 2283 2283 2283 Benzene 18647 18647 31 411 18206 18206 714 714 411 0 18206 0 0 0 3,3-Dimethylpentane 799 799 0 172 627 627 216 216 172 0 1835 1835 1835 1835 2-Methylhexane 4529 4529 0 1306 3223 3223 1606 1606 1306 0 9475 9475 9475 9475 2,3-Dimethylpentane 1865 1865 0 765 1100 1100 810 810 765 0 3241 3241 3241 3241 3-Ethylpentane 799 799 0 493 306 306 429 429 493 0 908 908 908 908 Table 1 -part 1(continued) STREAM 101 101a 102 103 104 104a 105 106 107 108 109 110 111 112 3-Methylhexane 5861 5861 0 2744 3117 3117 2819 2819 2744 0 9212 9212 9212 9212 T- 1 ,3-Dimethylcyclohexane 533 533 0 347 185 185 271 271 347 0 549 549 549 549 T- 1 ,2-Dimethylcycloxexane 533 533 0 350 183 183 271 271 350 0 543 543 543 543 N-Heptane 5328 5328 0 4267 1061 1061 3145 3145 4267 0 3183 3183 3183 3183 Methylcychexane 266 266 0 256 11 11 119 119 256 0 33 33 33 33 2-Methyl-1-Hexene 533 533 00 277 256 256 267 267 277 0 756 756 756 756 Ethylcyclopentane 799 799 00 767 32 32 362 362 767 0 97 97 97 97 2,5-Dimethylhexane 266 266 00 261 6 6 131 131 261 0 18 18 18 18 2,4-Dimethylhexane 533 533 00 524 9 9 250 250 524 0 27 27 27 27 Toluene 63934 63934 0 63511 423 423 2355723557 63511 0 1297 1297 1297 1297 2-Methyl heptane 1332 1332 0 1032 300 300 767 767 1032 0 902 902 902 902 3-Methylheptane 266 266 00 266 0 0 95 95 266 0 1 1 1 1 4-Methyl heptane 799 799 00 798 1 1 295 295 798 0 3 3 3 3 N -Octane 1332 1332 00 1332 0 0 413 413 1332 0 0 0 0 0 Ethyl benzene 9057 9057 0 9057 0 0 2263 2263 9057 0 0 0 0 0 M-xylene 2663926639 0 26639 0 0 6473 6473 26639 0 0 0 0 0 P-Xylene 10656 10656 0 10656 0 0 2602 2602 10656 0 0 0 0 0 O-Xylene 14652 14652 0 14652 0 0 3418 3418 14652 0 0 0 0 0 C-1 ,3-Dimethylcyclohexane 266 266 0 266 0 0 84 84 266 0 0 0 0 0 Nonane 533 533 0 533 0 0 123 123 533 0 0 0 0 0 Tetramethylbenzene 3729537295 0 37295 0 0 7023 702337295 0 0 0 0 0 Decane 6660 6660 0 6660 0 0 1355 1355 6660 0 0 0 0 0 Undecane 0 0 0 0 0 0 0 0 0 0 0 0 0 0 N-Pentylbenzene 0 0 0 0 0 0 0 0 0 0 0 0 0 0 Table 1 - part2 STREAM 113 114 115 116 117 118 119 120 121 122 Temperature F 268 268 201 212 204 254 105 105 268 234 Pressure psi 244 244 30 30 250 250 244 244 244 244 Mass Flow Ib/hr H2 17 12 0 0 0 0 1 232 5 6 Methane 388 272 0 0 0 0 64 1500 116 180 Ethane 1051 736 0 0 0 0 276 1299 315 590 Propane 1668 1169 0 0 0 0 486 723 499 986 I-Butane 31 22 28 28 23 23 8 5 9 17 N-Butane 1316 922 88 88 71 71 317 151 394 711 I-C5 2674 1874 1224 1224 991 991 468 100 801 1269 N-C5 2399 1681 1062 1062 860 860 381 63 718 1099 2,2 Dimethylbutane 894 626 475 475 389 389 109 12 268 377 2,3-Dimethylbutane 1935 1356 956 956 803 803 207 17 579 786 Cycpentane 0 0 0 0 0 0 0 0 0 0 Cycpentene 89 62 50 50 41 41 13 2 27 39 1-Pentene 49 34 30 30 25 25 8 2 15 23 2-Methylpentane 8653 6063 4179 4179 3569 3569 910 69 2590 3500 3-Methylpentane 11296 7915 5285 5285 4575 4575 1117 77 3381 4498 N-Hexane 18552 12999 8581 8581 7313 7313 1666 93 5553 7220 1-Hexene 476 333 223 223 193 193 48 3 142 190 Cychexene 1213 850 490 490 452 452 86 3 363 449 Cychexane 52730 36946 0 0 0 0 3688 149 15784 19472 2,2-Dimethylpentane 2119 1484 934 934 796 796 156 7 634 790 Methylcyclopentane 2735 1916 1244 1244 1060 1060 229 12 819 1047 2-4-Dimethylpentane 2123 1487 932 932 796 796 154 6 635 789 Benzene 0 0 20863 20863 18206 18206 0 0 0 0 3,3-Dimethylpentane 1723 1208 696 696 627 627 108 4 516 624 2-Methylhexane 8923 6252 3621 3621 3223 3223 537 15 2671 3208 2,3-Dimethylpentane 3056 2141 1262 1262 1100 1100 180 5 915 1095 3-Ethylpentane 860 602 385 385 306 306 47 1 257 304 Table 1 - part 2 (continued) STREAM 113 114 115 116 117 118 119 120 121 122 3-Methylhexane 8699 6095 3665 3665 3117 3117 500 13 2604 3103 T- 1 ,3-Dimethylcyclohexane 519 364 231 231 185 185 29 1 155 185 T- 1 ,2-Dimethylcycloxexane 514 360 229 229 183 183 29 1 154 183 N-Heptane 3029 2122 1569 1569 1061 1061 151 3 907 1058 Methylcychexane 31 22 19 19 11 11 1 0 9 11 2-Methyl-1-Hexene 715 501 306 306 256 256 41 1 214 255 Ethylcyclopentane 92 65 58 58 32 32 4 0 28 32 2,5-Dimethylhexane 17 12 12 12 6 6 1 0 5 6 2,4-Dimethylhexane 26 18 19 19 9 9 1 0 8 9 Toluene 1247 874 837 837 423 423 49 1 373 422 2-Methylheptane 859 602 439 439 300 300 42 1 257 299 3-Methylheptane 1 0 1 1 0 0 0 0 0 0 4-Methylheptane 3 2 2 2 1 1 0 0 1 1 N-Octane 0 0 0 0 0 0 0 0 0 0 Ethylbenzene 0 0 0 0 0 0 0 0 0 0 M-xylene 0 0 0 0 0 0 0 0 0 0 P-Xylene 0 0 0 0 0 0 0 0 0 0 O-Xylene 0 0 0 0 0 0 0 0 0 0 C-1 ,3-Dimethylcyclohexane 0 0 0 0 0 0 0 0 0 0 Nonane 0 0 0 0 0 0 0 0 0 0 Tetramethylbenzene 0 0 0 0 0 0 0 0 0 0 Decane 0 0 0 0 0 0 0 0 0 0 Undecane 0 0 0 0 0 0 0 0 0 0 N-Pentylbenzene 0 0 0 0 0 0 0 0 0 0 EXAMPLE 2 Structured Catalyst Packing The reactor was configured to behave as a down flow, plug flow reactor. This run used a single pass with 10% benzene in cyclohexane feed through the column filled with 1.09 lbs of dispersed nickel catalyst (KL-6564-T1.2 ) prepared in modules described in U.S. Pat. No.5,432, 890. The feed system was limited to a maximum flow rate of 16 Ib/hr. The differential pressure meter indicated nearly no pressure drop through the column. Stoichiometric hydrogen flow is 14.6 scf per Ib of benzene. During most of the run, the hydrogen to hydrocarbon feed ratio was kept at 2 scf H2 per Ib hydrocarbon, thus the hydrogen stoichiometry was 140%. The heat of reaction from the hydrogenation of one Ib of benzene is six times higher than the latent heat of vaporization for cyclohexane in a weight basis; thus the maximum concentration of benzene to prevent complete vaporization would have been 17%. Data from Run

There was a stronger correlation between conversion and temperature than there was between flow rate and conversion. Higher temperature resulted in a higher kinetic rate constant at the same flow rate but also a lower hydrogen partial pressure. The maximum WHSV obtainable to maintain 97% conversion was about 9. EXAMPLE 3 Dumped Catalyst The column was loaded with 4.5 lbs (14 feet) of the catalyst of Example 2 in a dumped bed. The feed flow was 40 Ib/hr of recycled cyclohexane and 10 lb/hr of 50% benzene in cyclohexane with 90 scfh of hydrogen. Complete conversion of the benzene to cyclohexane was obtained throughout the run. Hydrogen conversion was about 70%. Recycle can be used to increase the velocity through the reactor without having to make the reactor unusually long and thin. It also works to act as a heat sink for highly exothermic reactions, such as benzene hydrogenation, and will dilute the feed to a level that can be handled in the boiling point reactor described here. Recycle has limited applications because it dilutes the effectiveness of the reactor as well. Compared to a single pass, plug flow reactor, adding a recycle line where the recycle flow rate equals the feed flow rate (doubling the velocity), a kinetic improvement of 25% is required to achieve the same overall conversion. The structured catalyst packing (Example 2) provides lower pressure drops in the reactor. A structured bed with a void fraction of 0.5 will have 4 times less pressure drop compared to a dumped bed with a void fraction of 0.3, and a structured bed with a void fraction of 0.6 will have 8 times less pressure drop. However, with structured packing, conversion was 97% at a WHSV of 9, whereas with dumped packing, 99+ % conversion was obtained at a WHSV of 11 and higher. The apparent catalyst activity was three times higher in the dumped packing than in the structured packing, most likely due to improved mass transfer with higher flow velocity in the dumped packing, and increased bypassing in the dual void dimension of the structured packing. Preferably the down flow boiling point reactor is operated in a pulse flow mode to take advantage of the improved hydraulic mixing. The operating conditions for entering pulse flow based on generalized flow maps suggest, at 200 psig and 340 °F, the flow rates have to exceed 60 Ib/hrof hydrocarbon liquid and 90 scfh of vapor in order to enter the pulse flow hydrodynamic regime.