# Patent

## Title
Energy efficient method for exothermic reactions

## Abstract

An energy efficient process scheme for a highly exothermic reaction-distillation system in which the reactor is external to the distillation column and the feed to the reactor is a mixture of at least one liquid product stream from the distillation column with or without other liquid/vapor reactants. 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 pressure (and hence the temperature) is maintained at a sufficiently high level such that the reactor effluent can be efficiently used to provide reboil heat for the distillation column.

## Background
CROSS-REFERENCE TO RELATED APPLICATION
This application is a continuation application of U.S. patent application Ser. No. 11/038,755, now U.S. Pat. No. 7,320,745, filed Jan. 19, 2005.


BACKGROUND OF DISCLOSURE
1. Field of the Disclosure
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.
2. Background
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 DISCLOSURE
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.



BRIEF DESCRIPTION OF THE DRAWINGS 
The drawing is a schematic representation of a generalized implementation of the energy recovery process of the present invention.



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/O2/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 101a to the distillation column is a multi-component 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 101a. 
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. Under this 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 light-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 lb/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
223°
F.


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
lb/hr/ft2


Mass flow liquid out
5,683
lb/hr/ft2


Mass flow vapor in
1,832
lb/hr/ft2


Mass flow vapor out
7,000
lb/hr/ft2


Benzene conversion across reactor
99%


Reactor recycle mass ratio, stream 9/stream 15
2.36









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 streams 102 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 25







Draw stream locations











Stream 153
Stage
62



Stream 115
Stage
26



Stream 156
Stage
57



Stream 152
Stage
1 30



Stream 154
Stage
26







Reboiler duties











Reboiler 6 (external heat source)
23
MM btu/hr



Reboiler 7
9.8
MM btu/hr



Reboiler 8
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









STREAM















101
101a
102
103
104
104a
105





Temperature, F.
257
176
123
312
201
254
356


Pressure, psi
145
145
29
31
30
250
40


Mass Flow, lb/h


H2
0
0
0
0
0
0
0


Methane
0
0
0
0
0
0
0


Ethane
0
0
0
0
0
0
0


Propane
0
0
0
0
0
0
0


i-Butane
799
799
777
0
23
23
0


n-Butane
1865
1865
1794
0
71
71
0


i-C5
117
117
10730
0
991
991
0


n-C5
8258
8258
7398
0
860
860
0


2,2-Dimethylbutane
1865
1865
1476
0
389
389
0


2,3-Dimethylbutane
2131
2131
1328
0
803
803
0


Cyclopentane
0
0
0
0
0
0
0


Cyclopentene
266
266
226
0
41
41
0


1-Pentene
266
266
242
0
25
25
0


2-Methylpentane
7459
7459
3890
0
3569
3569
0


3-Methylpentane
6127
6127
1552
0
4575
4575
0


n-Hexane
7459
7459
146
0
7313
7313
0


1-Hexene
266
266
73
0
193
193
0


Cyclohexene
533
533
0
81
452
452
12


Cyclohexane
0
0
0
0
0
0
0


2,2-Dimethylpentane
799
799
0
0
796
796
7


Methylcyclopentane
1066
1066
5
1
1060
1060
3


2,4-Dimethylpentane
799
799
0
0
796
796
9


Benzene
18647
18647
31
411
18206
18206
714


3,3-Dimethylpentane
799
799
0
172
627
627
216


2-Methylhexane
4529
4529
0
1306
3223
3223
1606


2,3-Dimethylpentane
1865
1865
0
765
1100
1100
810


3-Ethylpentane
799
799
0
493
306
306
429


3-Methylhexane
5861
5861
0
2744
3117
3117
2819


t-1,3-Dimethylcyclohexane
533
533
0
347
185
185
271


t-1,2-Dimethylcyclohexane
533
533
0
350
183
183
271


n-Heptane
5328
5328
0
4267
1061
1061
3145


Methylcyclohexane
266
266
0
256
11
11
119


2-Methyl-1-Hexene
533
533
0
277
256
256
267


Ethylcyclopentane
799
799
0
767
32
32
362


2,5-Dimethylhexane
266
266
0
261
6
6
131


2,4-Dimethylhexane
533
533
0
524
9
9
250


Toluene
63934
63934
0
63511
423
423
23557


2-Methylheptane
1332
1332
0
1032
300
300
767


3-Methylheptane
266
266
0
266
0
0
95


4-Methylheptane
799
799
0
798
1
1
295


n-Octane
1332
1332
0
1332
0
0
413


Ethylbenzene
9057
9057
0
9057
0
0
2263


m-Xylene
26639
26639
0
26639
0
0
6473


p-Xylene
10656
10656
0
10656
0
0
2602


o-Xylene
14652
14652
0
14652
0
0
3418


o-1,3-Dimethylcyclohexane
266
266
0
266
0
0
84


Nonane
533
533
0
533
0
0
123


Tetramethylbenzene
37295
37295
0
37295
0
0
7023


Decane
6660
6660
0
6660
0
0
1355


Undecane
0
0
0
0
0
0
0


n-Pentylbenzene
0
0
0
0
0
0
0












STREAM















106
107
108
109
110
111
112





Temperature, F.
290
194
86
223
359
320
268


Pressure, psi
31
31
250
250
245
245
245


Mass Flow, lb/h


H2
0
0
1649
1660
251
251
251


Methane
0
0
1680
1952
1952
1952
1952


Ethane
0
0
1889
2626
2626
2626
2626


Propane
0
0
1709
2878
2878
2878
2878


i-Butane
0
0
0
44
44
44
44


n-Butane
0
0
791
1784
1784
1784
1784


i-C5
0
0
378
3243
3243
3243
3243


n-C5
0
0
302
2843
2843
2843
2843


2,2-Dimethylbutane
0
0
0
1015
1015
1015
1015


2,3-Dimethylbutane
0
0
0
2159
2159
2159
2159


Cyclopentane
0
0
0
0
0
0
0


Cyclopentene
0
0
0
103
103
103
103


1-Pentene
0
0
0
59
59
59
59


2-Methylpentane
0
0
0
9632
9832
9832
9832


3-Methylpentane
0
0
0
12489
12489
12489
12489


n-Hexane
0
0
0
20312
20312
20312
20312


1-Hexene
0
0
0
526
526
526
526


Cyclohexene
12
81
0
1302
1302
1302
1302


Cyclohexane
0
0
0
36946
56567
56567
56567


2,2-Dimethylpentane
7
0
0
2281
2281
2281
2281


Methylcyclopentane
3
1
0
2976
2976
2976
2976


2,4-Dimethylpentane
9
0
0
2283
2283
2283
2283


Benzene
714
411
0
18206
0
0
0


3,3-Dimethylpentane
216
172
0
1835
1835
1835
1835


2-Methylhexane
1606
1306
0
9475
9475
9475
9475


2,3-Dimethylpentane
810
765
0
3241
3241
3241
3241


3-Ethylpentane
429
493
0
908
908
908
908


3-Methylhexane
2819
2744
0
9212
9212
9212
9212


t-1,3-Dimethylcyclohexane
271
347
0
549
549
549
549


t-1,2-Dimethylcyclohexane
271
350
0
543
543
543
543


n-Heptane
3145
4267
0
3183
3183
3183
3183


Methylcyclohexane
119
256
0
33
33
33
33


2-Methyl-1-Hexene
267
277
0
756
756
756
756


Ethylcyclopentane
362
767
0
97
97
97
97


2,5-Dimethylhexane
131
261
0
18
18
18
18


2,4-Dimethylhexane
250
524
0
27
27
27
27


Toluene
23557
63511
0
1297
1297
1297
1297


2-Methylheptane
767
1032
0
902
902
902
902


3-Methylheptane
95
266
0
1
1
1
1


4-Methylheptane
295
798
0
3
3
3
3


n-Octane
413
1332
0
0
0
0
0


Ethylbenzene
2263
9057
0
0
0
0
0


m-Xylene
6473
26639
0
0
0
0
0


p-Xylene
2602
10656
0
0
0
0
0


o-Xylene
3418
14652
0
0
0
0
0


o-1,3-Dimethylcyclohexane
84
266
0
0
0
0
0


Nonane
123
533
0
0
0
0
0


Tetramethylbenzene
7023
37295
0
0
0
0
0


Decane
1355
6660
0
0
0
0
0


Undecane
0
0
0
0
0
0
0


n-Pentylbenzene
0
0
0
0
0
0
0












STREAM













113
114
115
116
117





Temperature, F.
268
268
201
212
204


Pressure, psi
244
244
30
30
250


Mass Flow, lb/h


H2
17
12
0
0
0


Methane
388
272
0
0
0


Ethane
1051
736
0
0
0


Propane
1668
1169
0
0
0


i-Butane
31
22
28
28
23


n-Butane
1316
922
88
88
71


i-C5
2674
1874
1224
1224
991


n-C5
2399
1681
1062
1062
860


2,2-Dimethylbutane
894
626
475
475
389


2,3-Dimethylbutane
1935
1356
956
956
803


Cyclopentane
0
0
0
0
0


Cyclopentene
89
62
50
50
41


1-Pentene
49
34
30
30
25


2-Methylpentane
8653
6063
4179
4179
3569


3-Methylpentane
11296
7915
5285
5285
4575


n-Hexane
18552
12999
8581
8581
7313


1-Hexene
476
333
223
223
193


Cyclohexene
1213
850
490
490
452


Cyclohexane
52730
36946
0
0
0


2,2-Dimethylpentane
2119
1484
934
934
796


Methylcyclopentane
2735
1916
1244
1244
1060


2,4-Dimethylpentane
2123
1487
932
932
796


Benzene
0
0
20863
20863
18206


3,3-Dimethylpentane
1723
1208
696
696
627


2-Methylhexane
8923
6252
3621
3621
3223


2,3-Dimethylpentane
3056
2141
1262
1262
1100


3-Ethylpentane
860
602
385
385
306


3-Methylhexane
8699
6095
3665
3665
3117


t-1,3-Dimethylcyclohexane
519
364
231
231
185


t-1,2-Dimethylcyclohexane
514
360
229
229
183


n-Heptane
3029
2122
1569
1569
1061


Methylcyclohexane
31
22
19
19
11


2-Methyl-1-Hexene
715
501
306
306
256


Ethylcyclopentane
92
65
58
58
32


2,5-Dimethylhexane
17
12
12
12
6


2,4-Dimethylhexane
26
18
19
19
9


Toluene
1247
874
837
837
423


2-Methylheptane
859
602
439
439
300


3-Methylheptane
1
0
1
1
0


4-Methylheptane
3
2
2
2
1


n-Octane
0
0
0
0
0


Ethylbenzene
0
0
0
0
0


m-Xylene
0
0
0
0
0


p-Xylene
0
0
0
0
0


o-Xylene
0
0
0
0
0


o-1,3-Dimethylcyclohexane
0
0
0
0
0


Nonane
0
0
0
0
0


Tetramethylbenzene
0
0
0
0
0


Decane
0
0
0
0
0


Undecane
0
0
0
0
0


n-Pentylbenzene
0
0
0
0
0












STREAM















118
119
120
121
122







Temperature, F.
254
105
105
268
234



Pressure, psi
250
244
244
244
244



Mass Flow, lb/h



H2
0
1
232
5
6



Methane
0
64
1500
116
180



Ethane
0
276
1299
315
590



Propane
0
486
723
499
986



i-Butane
23
8
5
9
17



n-Butane
71
317
151
394
711



i-C5
991
468
100
801
1269



n-C5
860
381
63
718
1099



2,2-Dimethylbutane
389
109
12
268
377



2,3-Dimethylbutane
803
207
17
579
786



Cyclopentane
0
0
0
0
0



Cyclopentene
41
13
2
27
39



1-Pentene
25
8
2
15
23



2-Methylpentane
3569
910
69
2590
3500



3-Methylpentane
4575
1117
77
3381
4498



n-Hexane
7313
1666
93
5553
7220



1-Hexene
193
48
3
142
190



Cyclohexene
452
86
3
363
449



Cyclohexane
0
3688
149
15784
19472



2,2-Dimethylpentane
796
156
7
634
790



Methylcyclopentane
1060
229
12
819
1047



2,4-Dimethylpentane
796
154
6
635
789



Benzene
18206
0
0
0
0



3,3-Dimethylpentane
627
108
4
516
624



2-Methylhexane
3223
537
15
2671
3208



2,3-Dimethylpentane
1100
180
5
915
1095



3-Ethylpentane
306
47
1
257
304



3-Methylhexane
3117
500
13
2604
3103



t-1,3-Dimethylcyclohexane
185
29
1
155
185



t-1,2-Dimethylcyclohexane
183
29
1
154
183



n-Heptane
1061
151
3
907
1058



Methylcyclohexane
11
1
0
9
11



2-Methyl-1-Hexene
256
41
1
214
255



Ethylcyclopentane
32
4
0
28
32



2,5-Dimethylhexane
6
1
0
5
6



2,4-Dimethylhexane
9
1
0
8
9



Toluene
423
49
1
373
422



2-Methylheptane
300
42
1
257
299



3-Methylheptane
0
0
0
0
0



4-Methylheptane
1
0
0
1
1



n-Octane
0
0
0
0
0



Ethylbenzene
0
0
0
0
0



m-Xylene
0
0
0
0
0



p-Xylene
0
0
0
0
0



o-Xylene
0
0
0
0
0



o-1,3-Dimethylcyclohexane
0
0
0
0
0



Nonane
0
0
0
0
0



Tetramethylbenzene
0
0
0
0
0



Decane
0
0
0
0
0



Undecane
0
0
0
0
0



n-Pentylbenzene
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. 5,431,890. The feed system was limited to a maximum flow rate of 16 lb/hr. The differential pressure meter indicated nearly no pressure drop through the column.
Stoichiometric hydrogen flow is 14.6 scf per lb of benzene. During most of the run, the hydrogen to hydrocarbon feed ratio was kept at 2 sef H2 per lb hydrocarbon, thus the hydrogen stoichiometry was 140%. The heat of reaction from the hydrogenation of one lb 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















HC Rate
H2 Rate
Inlet T
Average T
Exit T
Bz Conv
H2 Conv
WHSV
Index


(lb/h)
(scfh)
(° F.)
(° F.)
(° F.)
(%)
(%)
(h−1)
(psia−1h−1)


















2
8
250
320
300
99.8+
34
1.8
0.10


10
20
200
326
350
97
70
9.2
0.30


16
32
200
280
340
89
60
15
0.2









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 lbs/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 lb/hr of hydrocarbon liquid and 90 scfh of vapor in order to enter the pulse flow hydrodynamic regime.

## Claims

1. A 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 the multi component petroleum to recover an overhead fraction, a bottoms fraction, and at least two intermediate fractions including a first intermediate fraction containing the reaction component having a first boiling range, and a second intermediate fraction having a second boiling range which is lower than the first boiling range;
removing the first intermediate fraction from the fractional distillation;
reacting the reaction component under conditions to exothermically react the reaction component and produce a reaction stream having a temperature higher than the second boiling range;
removing the second intermediate fraction from the fractional distillation;
heating the second intermediate fraction by indirect heat exchange with the reaction stream; and
returning the heated second intermediate fraction to the fractional distillation.

2. The process according to claim 1 further comprising:
recovering a third intermediate fraction having a third boiling range which is lower than the first boiling range;
removing the third intermediate fraction from the fractional distillation;
heating the third intermediate fraction by indirect heat exchange with the reaction stream, and;
returning the heated third intermediate fraction to the fractional distillation.

3. The process of claim 2, wherein the second boiling range is lower than the third boiling range, the process comprising heating the third intermediate fraction prior to heating the second intermediate fraction.

4. The process of claim 1, wherein the reaction component comprises benzene and the reacting comprises forming at least one of ethylbenzene, cumene, and butylbenzene.

5. The process of claim 1, wherein the reaction component comprises nitrobenzene and the reacting comprises forming aniline.

6. The process of claim 1, wherein the reaction component comprises carbon monoxide and the reacting comprises forming methanol.

7. The process of claim 1, wherein the reaction component comprises at least one of acetic acid and ethylene, and the reacting comprises forming vinylacetate.

8. The process of claim 1, wherein the reacting comprises operating a fixed bed downflow reactor at a boiling point of a mixture comprising the first intermediate fraction and a reaction product.

9. The process of claim 8, wherein the reacting comprises operating the fixed bed downflow reactor in a pulse flow mode.