HYDROCARBON STREAMS BACKGROUND OF THE INVENTION Field of the Invention

The present invention relates to the selective hydrogenation of diolefins and acetylenic compounds and isomerization of olefins to more desirable isomers in an olefin rich stream containing sulfur impurities. More particularly the invention relates to a process utilizing hydrogenation catalysts in a structure to serve as both the catalyst and as a distillation structure for the simultaneous reaction and separation of the reactants and reaction products. The present process includes the removal of mercaptans, hydrogen sulfide (H ₂ S) and polyolefins from petroleum distillate streams. More particularly the invention relates to a process wherein the petroleum distillate contains diolefins which are selectively reacted with the mercaptans and/or hydrogen sulfide (H ₂ S) to form sulfides and the remaining diolefins and acetylenes are hydrogenated to mono-olefins. More particularly the invention relates to a process wherein the reaction of the mercaptans and/or hydrogen sulfide (H ₂ S) with the diolefins is carried out simultaneously with a fractional distillation to remove the sulfides, and thus the sulfur, from the distillate. Most particularly the invention relates to a process wherein most of the hydrogen sulfide is removed prior to the reaction of the diolefins and mercaptans.

Related Information

Petroleum distillate streams contain a variety of organic chemical components. Generally the streams are defined by their boiling ranges which determine the compositions. The processing of the streams also affects the composition. For instance, products from either catalytic cracking or thermal cracking processes contain high concentrations of olefinic materials as well as

saturated (alkanes) materials and polyunsaturated materials (diolefins) . Additionally, these components may be any of the various isomers of the compounds.

The petroleum distillates often contain unwanted contaminants such as sulfur and nitrogen compounds. These contaminants often are catalyst poisons or produce undesirable products upon further processing. In particular the sulfur compounds can be troublesome. The sulfur compounds are known catalyst poisons for naphtha reforming catalysts and hydrogenation catalysts. The sulfur compounds present in a stream are dependent upon the boiling range of the distillate. In a light naphtha (110- 420°F boiling range) the predominant sulfur compounds are mercaptans. Streams having C ₃ hydrocarbons also may contain H ₂ S. The most common method for removal of the H ₂ S is amine extraction and the most common method for removal of mercaptans is caustic washing of the organic streams.

Another method of removal of the sulfur compounds is by hydrodesulfurization (HDS) in which the petroleum distillate is passed over a solid particulate catalyst comprising a hydrogenation metal supported on an alumina base. Additionally copious quantities of hydrogen are included in the feed. The following equations illustrate the reactions in a typical HDS unit:

(1) RSH + H ₂ ► RH + H ₂ S

(2) RCI + H ₂ ► RH + HCl

(3) 2RN + 4H ₂ ► RH + NH ₃

(4) ROOH + 2H ₂ ► RH + H ₂ 0 Typical operating conditions for the HDS reactions are:

Temperature, °F 600-780

Pressure, psig 600-3000

H ₂ recycle rate, SCF/bbl 1500-3000 Fresh H ₂ makeup, SCF/bbl 700-1000

As may be seen the emphasis has been upon hydrogenating the sulfur and other contaminating compounds. The sulfur is

then removed in the form of gaseous H ₂ S, which in itself is a pollutant and requires further treatment.

In addition to sulfur and nitrogen compounds mixed refinery streams contain a broad spectrum of olefinic compounds. This is especially true of products from either catalytic cracking or thermal cracking processes. These unsaturated compounds comprise ethylene, acetylene, propylene, propadiene, methyl acetylene, butenes, butadiene, amylenes, hexenes etc. Many of these compounds are valuable, especially as feed stocks for chemical products. Ethylene, especially is recovered. Additionally, propylene and the butenes are valuable. However, the olefins having more than one double bond and the acetylenic compounds (having a triple bond) have lesser uses and are detrimental to many of the chemical processes in which the single double bond compounds are used, for example polymerization. Over the range of hydrocarbons under consideration, the removal of highly unsaturated compounds is of value as a feed pretreatment, since these compounds have frequently been found to be detrimental in most processing, storage and use of the streams.

In the production of tertiary amyl methyl ether (TAME) for use as a gasoline additive generally a light cracked naphtha (LCN) is used as the source of the olefins for the etherification reaction. The acetylenes and diolefins are detrimental in the etherification process as well as in other processes such as alkyiation and should be removed early in the stream processing. The LCN usually contains sulfur as a contaminant in the form of mercaptans in concentrations of up to hundreds wppm. These mercaptans are inhibitors for the hydrogenation catalyst used to hydrogenate dienes and acetylenes and obtain beneficial isomerization in the feed to an etherification unit or to an alkyiation unit. Although the most desirable hydrogenation catalysts are inhibited by sulfur compounds even in very small amounts, e.g. 10-100 ppm, there are other similar catalysts that will cause the sulfur compounds and the diolefins to form

adducts, which can be separated from the lighter components.

It is an advantage of the present invention that the sulfur compounds can be separated from the lighter hydrocarbon components which can then be hydrotreated with the sulfur sensitive catalyst to hydrogenate highly unsaturated hydrocarbons and obtain beneficial isomerization of the mono-olefins. It is a particular advantage that this may be achieved in a single reactive distillation column by using beds of function specific catalyst. It is a particular feature of the present invention that a dual bed system may be used.

SUMMARY OF THE INVENTION The present invention presents a new process for the removal of mercaptans and/or hydrogen sulfide (H ₂ S) from aliphatic hydrocarbon streams, containing 3 to 12 carbon atoms comprising distilling a hydrocarbon stream comprising C ₃ to C ₁₂ hydrocarbons including alkanes, mono- olefins, diolefins, acetylenes and minor amounts of sulfur compounds to remove a fraction comprising at least C ₃ 's and a portion of the sulfur compounds and leaving a residual, concurrently:

(1) feeding hydrogen and a portion of the residual to distillation column reactor containing a first bed comprising a first hydrogenation catalyst of the type characterized by nickel, cobalt or iron, preferably selected from nickel, cobalt, iron or mixtures thereof, and prepared in the form of a distillation structure and a second bed position in said column above said first bed, said second bed comprising a second hydrogenation catalyst of the type characterized by platinum, palladium or rhodium, preferably selected from platinum, palladium, rhodium or mixtures thereof, and prepared as a distillation structure wherein any sulfur compounds in the residual react in said first bed with a portion of the diolefins to form sulfides in a first reaction mixture,

(2) fractionally distilling the first reaction

mixture to remove the sulfides with a heavier fraction and passing a lighter fraction into the second bed

(3) hydrogenating the diolefins and acetylenes in said second bed to form a second reaction mixture (4) fractionally distilling the second reaction mixture and

(5) removing a fraction overhead, which is substantially free of sulfur compounds, acetylenes and diolefins. In step 3 the hydrogenation of the more highly unsaturated compounds will produce more mono-olefins and/or alkanes. Also there may be bond shifting isomerization such as butene-2 to butene-l.

In the first (lower) bed the catalytic material may be initially present as the oxide or as reduced metal but is converted to the sulfide form during the reaction.

Generally the catalytic material in the second (upper) bed is initially present as the metal oxide and may be converted to the hydride form during use by the hydrogen. In the present invention hydrogen is provided at an effectuating hydrogen partial pressure of at least about

0.1 psia to less than 70 psia, preferably less than 50 psia, more preferably less than 35 psia to the distillation column reactor containing hydrogenation catalysts as described.

BRIEF DESCRIPTION OF THE DRAWING The FIGURE is a flow diagram in εchematic form of a preferred embodiment of the invention.

DESCRIPTION OF THE PREFERRED EMBODIMENTS The preεent invention provides a process for the reaction of diolefins and acetylenes within a petroleum distillate with the mercaptans within the distillate to form sulfides and concurrent separation of the higher boiling sulfides from the distillate; selectively hydrogenating remaining diolefins and acetylenes and isomerizing the mono-olefins to equilibrium. This requires a distillation column reactor which contains at leaεt two beds of an appropriate catalyst in the form of a

catalytic distillation structure.

In the usual application of a process where the catalyst serves as a distillation component, the equilibrium is constantly disturbed, thus driving the reaction toward completion, that is, the reaction has an increased driving force because the reaction products have been removed and cannot contribute to a reverse reaction (LeChatelier's Principle) . Although the hydrogenation reactions have been described as reversible at elevated temperatures above about 900°F, under the temperature conditions employed in the present invention, the hydrogenation is not reversible and cannot be an incentive to use a catalytic distillation system. The poor performance of prior vapor phase hydrogenations would not suggest the use of distillation type reaction.

It is believed that in the present reaction catalytic distillation is a benefit first, because the reaction is occurring concurrently with distillation, the initial reaction products and other stream components are removed from the reaction zone as quickly as posεible reducing the likelihood of εide reactionε. Second, because all the components are boiling the temperature of reaction is controlled by the boiling point of the mixture at the system pressure. The heat of reaction simply creates more boil up, but no increase in temperature at a given pressure. As a result, a great deal of control over the rate of reaction and distribution of products can be achieved by regulating the system pressure. Also, adjusting the throughput (residence time = liquid hourly space velocity" ¹ ) gives further control of product distribution and to a degree control of the side reactions such as oligomerization. A further benefit that this reaction may gain from catalytic distillation is the washing effect that the internal reflux provides to the catalyst thereby reducing polymer build up and coking. Internal reflux may be varied over the range of 0.2 to 20 L/D (wt. liquid just below the catalyst bed/wt. distillate) and gives excellent results, and with the C ₃ -C ₅ streams

being usually in the range of 0.5 to 4 L/D.

Quite surprisingly the low hydrogen partial pressure used in the distillation εystem did not result in the failure of the hydrogenation which would have been expected based on the high hydrogen partial pressure found in the liquid phase systems which are the worldwide standard. Without limiting the scope of the invention it is proposed that the mechanism that produces the effectiveness of the present process is the condensation of a portion of the vapors in the reaction system, which occludes sufficient hydrogen in the condensed liquid to obtain the requisite intimate contact between the hydrogen and the highly unsaturated compounds in the presence of the catalyst to result in their hydrogenation. This phenomenon of condensation which is a constant factor in a distillation is believed to result in the same or better hydrogen availability, as the high pressure in the liquid phase, that is, the hydrogen is introduced into the liquid so that the hydrogenation occurs. The C ₅ 's in the feed to the present unit are contained in a single "light naphtha" cut which may contain everything from C ₃ 's through Cg's and higher. This mixture can easily contain 150 to 200 components. Mixed refinery streams often contain a broad εpectrum of olefinic compounds. This is especially true of products from either catalytic cracking or thermal cracking procesεeε. Refinery streams are usually separated by fractional distillation, and because they often contain compounds that are very close in boiling points, such separations are not precise. A C5 stream, for instance, may contain C ₃ 's and up to Cg's. These components may be saturated (alkanes) , unsaturated (mono-olefins) , or poly-unεaturated (diolefins) . Additionally, the components may be any or all of the various isomers of the individual compounds. Such streams typically contain 15 to 30 weight % of the isoamylenes.

Such refinery streamε alεo contain εmall amounts of sulfur which must be removed. The sulfur compounds are generally found in a light cracked naphtha stream aε

mercaptans and/or hydrogen sulfide (H ₂ S) which inhibit the hydrogenation catalyst used to selectively hydrogenate diolefins. Removal of sulfur compounds is generally termed "sweetening" a stream. 5 Several of the minor components (diolefins) in the feed will react slowly with oxygen during storage to produce "gum" and other undesirable materials. However, these components also react very rapidly in catalytic etherifications to form a yellow, foul smelling gummy

10 material and consume acid in an alkyiation unit. Thus it is seen to be desirable to remove these components whether the "light naphtha" cut is to be used only for gasoline blending by itself or as feed to a TAME (tertiary amyl methyl ether) or alkyiation procesε.

15. Catalyεtε which are uεeful in all the reactionε include the Group VIII metalε. The preferred catalyst for the ercaptan-diolefin reaction (lower bed) is nickel. The preferred catalyst for the εelective hydrogenation and iεomerization (upper bed) is palladium. The palladium

20 catalyst is inhibited by the presence of sulfur compounds and is thus placed above the nickel catalyεt in the diεtillation column reactor εuch that the feed iε firεt εubjected to the nickel catalyεt and the sulfur compounds removed by forming adducts with a portion of the diolefinε

25 in the feed. The catalyst may be uεe aε individual Group VIII metal componentε or in admixture with each other or modifierε aε known in the art, particularly those in Group VIB and IB.

Generally the metals are deposited as the oxideε on an 0 alumina εupport. The εupportε are usually small diameter extrudates or εphereε, typically alumina. The catalyεt muεt then be prepared in the form of a catalytic diεtillation εtructure. The catalytic distillation structure must be able to function as catalyst and aε maεs 5 transfer medium. The catalyst must be suitably supported and spaced within the column to act as a catalytic distillation structure. In a preferred embodiment the catalyst is contained in a woven wire mesh structure as

disclosed in U.S. Patent No. 5,266,546 which is hereby incorporated by reference. Other catalytic distillation structures useful for this purpose are disclosed in U.S. patents 4,73 ₁ ,229, 5,073,236 and U.S. Serial No. 08/188,803 filed ₀ 1/ ₃₁ /94 which are also incorporated by reference.

In a preferred embodiment a light cracked stream which is use _d as a feed to an etherification or alkyiation unit is the feed for this process. The light cracked naphtha contains _C3 's to Cg's components which may be saturated ₍ alkanes) , unsaturated (olefins) and poly-unsaturated (diolefins) along with minor amounts of the mercaptans. The light naphtha is generally depentanized in a fractional _d iεtillation column to remove that portion containing the _C6 an _d higher boiling materialε (C ₆ +) as bottoms and the C ₅ and lower boiling materials (C ₅ -) as overheads.

In the present invention the εtream iε first subjected to a distillation in a diεtillation vessel wherein the _C3 's, H _2S an _d a portion of the C ₄ 'ε are diεtilled overhead. The _C3 'ε over _h eadε may be εubjected to the traditional amine extraction to remove the H ₂ S. The bottoms from the _d istillation (C ₄ +) are fed to a debutanizer or _d epentanizer which has two separate distillation reaction _b e _d s in the rectification εection.

The lower bed contains a nickel sulfide catalytic _d istillation component to first react subεtantially all of the mercaptanε (and reεidual H ₂ S) contained in the light cracke _d nap _h tha with a portion of the diolefins to form sulfi _d eε which are higher boiling than the C ₅ fraction containing the amyleneε which are fed to the etherification and/or alkyiation unit. The sulfides are removed as _b ottomε from the depentanizer along with the C ₆ + fraction an _d can _b e remixed into the final gaεoline fraction.

The upper bed containε a εupported palladium catalytic diεtillation εtructure which selectively hydrogenates the remainder of the diolefins while at the same time isomerizing the mono-olefinε to equilibrium.

Hy _d rogen is provided as neceεεary to support the reaction. The distillation column reactor is operated at a

pressure such that the reaction mixture is boiling in the bed of catalyst. A "froth level", as described in U.S. Pat. No. 5,221,441 which is incorporated herein, may be maintained throughout the catalyst bed by control of the bottoms and/or overheads withdrawal rate, although the preferred operation is without the froth. As may be appreciated in the froth mode the liquid is boiling and the physical εtate is actually a froth having a higher density than would be normal in a packed distillation column but less than the liquid without the boiling vapors.

The present process preferably operates at overhead pressure of said distillation column reactor in the range between 0 and 250 psig and temperatures within said distillation reaction zone in the range of 100 to 300°F, preferably 130 to 270°F.

The feed and the hydrogen are preferably fed to the distillation column reactor separately or they may be mixed prior to feeding. A mixed feed is fed below the lower catalyst bed or at the lower end of the bed. Hydrogen alone is fed below the catalyst bed and the hydrocarbon stream is fed below the first bed to about the mid one- third of the first bed. The pressure selected is that which maintains catalyst bed temperature between 100°F and 300°F. A preferred catalyεt for the mercaptan-diolefin reaction iε 54 wt% Ni on 8 to 14 mesh alumina sphereε, supplied by Calcicat, designated as E-475-SR. Typical physical and chemical properties of the catalyst as provided by the manufacturer are as follows: TABLE I

Designation E-475-SR

Form Spheres

Nominal εize 8x14 Mesh

Ni wt% 54 Support Alumina

A p r e f e rre d c at a l y s t f o r the s e l e ct ive hydrogenation/isomerization reactions is palladium oxide, preferably 0. 1 to 5 . 0 weight % , supported on an

appropriate support medium such as alumina, carbon or silica, e.g., 1/8" alumina extrudates. The catalyst used is 0.4 wt% Pd on 1/8" Al ₂ 0 ₃ (alumina) extrudates, hydrogenation catalyst, supplied by United Catalysts, Inc. designated as G68C-1. Typical physical and chemical properties of the catalyst as provided by the manufacturer are as follows:

TABLE II Deεignation G68C-1 Form εphereε

Nominal size 8x12 Mesh

Pd. wt% 0.4

Support High purity alumina

The hydrogen rate to the distillation column reactor must be sufficient to maintain the reaction, but kept below that which would cause flooding of the column which is understood to be the "effectuating amount of hydrogen" as that term is used herein. Generally the mole ratio of hydrogen to diolefins and acetylenes in the feed is at leaεt 1.0 to 1.0, preferably at least 2.0 to 1.0 and more preferably at leaεt 10 to 1.0.

The nickel catalyεt alεo catalyzeε the selective hydrogenation of the diolefins contained within the light cracked naphtha and to a lesser degree the isomerization of some of the mono-olefins. However, the palladium catalyst iε preferred for these reactions. Generally the relative absorption preference iε aε follows:

(1) sulfur compounds

(2) diolefins (3) ono-olefinε

If the catalyεt sites are occupied by a more strongly abεorbed εpecies, reaction of these weaker abεorbed εpecies cannot occur. For this reason the εulfur compoundε are removed utilizing the nickel catalyεt. The reaction of intereεt in the nickel catalyεt bed iε the reaction of the mercaptanε and, to a lesser extent hydrogen sulfide (H ₂ S) with diolefins. The equation of intereεt which describes the reaction is:

H ₂ I

RSH + R ₁ C=C-C=C-R ₂ ^► R-S-C-C-C=R ₂ Ni I

H Where R, R^ and R are independently selected from hydrogen and hydrocarbyl groups of 1 to 20 carbon atoms. If there is concurrent hydrogenation of the dienes, then hydrogen will be consumed in that reaction.

Typical of the mercaptan compounds which may be found to a greater or lesser degree in a light cracked naphtha are: methyl mercaptan (b.p. 43 ^β F), ethyl mercaptan (b.p. 99°F), n-propyl mercaptan (b.p. 154°F), iso-propyl mercaptan (b.p. 135-140°F), iso-butyl mercaptan (b.p. 190°F), tert-butyl mercaptan (b.p. 147°F), n-butyl mercaptan (b.p. 208°F), sec-butyl mercaptan (b.p. 203°F), iso-amyl mercaptan (b.p. 250 ^β F) , n-amyl mercaptan (b.p. 259"F), α-methylbutyl mercaptan (b.p. 234 ^β F), α-ethylpropyl mercaptan (b.p. 293°F), n-hexyl mercaptan (b.p. 304°F), 2-mercapto hexane (b.p. 284°F), and 3-mercapto hexane (b.p. 135°F at 20 mm Hg) . The reaction of H ₂ S with the diolefins has been found to be considerably slower than the other sulfideε and thuε the preferred process removes the H S prior to subjecting the feed to the two catalyst beds.

The reactions of the C ₄ 's of interest are:

(1) butadiene-1,3 + hydrogen to butene-1 and butene-2 and

(2) butene-2 to butene-1. The reactions of the C ₅ ^! ε of intereεt are:

(1) iεoprene (2-methyl butadiene-1,3) + hydrogen to 2- methyl butene-1, 2-methyl butene-2 and 3-methyl butene-1;

(2) ciε- and tranε 1,3-pentadieneε (cis and trans piperylenes) + hydrogen to pentene-1 and pentene-2; (3) 3-methyl butene-1 to 2-methyl butene-2 and 2-methyl butene-1;

(4) 2-methyl butene-1 to 2-methyl butene-2; and

(5) 2-methyl butene-2 to 2-methyl butene-1.

The first two C ₅ reactions remove the undesirable components while the third is advantageous for feed to a

TAME reactor. The 3-methyl butene-1 does not react with methanol to produce TAME over the sulfonic acid catalyst while the two 2-methyl butenes do.

The present invention carries out the method in a catalyst packed column which can be appreciated to contain a vapor phase and εome liquid phase as in any diεtillation. The diεtillation column reactor iε operated at a pressure such that the reaction mixture iε boiling in the bed of catalyst. The present process operates at overhead presεure of εaid diεtillation column reactor in the range between 0 and 350 pεig, preferably 250 or leεs and temperatures within said distillation reaction zone in the range of 40 to 300°F, preferably 110 to 270°F at the requisite hydrogen partial pressures. The feed weight hourly space velocity (WHSV) , which iε herein understood to mean the unit weight of feed per hour entering the reaction distillation column per unit weight of catalyst in the catalytic distillation structures, may vary over a very wide range within the other condition perimeters, e.g. 0.5 to 35.

The advantages of utilizing a distillation column reactor in the instant selective hydrogenation process lie in the better selectivity of diolefin to olefin, conservation of heat and the separation by distillation which can remove some undesirable compounds, e.g. the sulfur contaminants, from the feed prior to expoεure to the hydrogenation/isomerization catalyst (the sulfides which are produced in the lower nickel catalyst bed are higher boiling than the C ₄ 's and Cs's so are distilled downward in the column away from the upper palladium catalyst bed) and the distillation can concentrate desired components in the catalyst zone.

The temperature in the distillation column reactor is determined by the boiling point of the liquid mixture present at any given presεure. The temperature in the lower portions of the column will reflect the constitution of the material in that part of the column, which will be higher than the overhead; that is, at constant presεure a

change in the temperature of the syεtem indicates a change in the composition in the column. To change the temperature the pressure is changed. Temperature control in the reaction zone is thus effected by a change in presεure; by increaεing the preεεure, the temperature in the εyεtem iε increased, and vice versa.

Referring now to the FIGURE there is depicted a simplified flow diagram of one embodiment of the invention. The combined C ₃ -naphtha εtream containing the olefinε, diolefins, mercaptans, and H ₂ S is first fed via flow line 100 to a distillation column 10 where the C ₃ 's and subεtantially all of the H S is distilled overhead and removed via flow line 102 for further sweetening as necesεary. Generally a portion of the C ₄ 's muεt be included in the overheadε to insure that essentially all of the H ₂ S is removed. The bottomε from the distillation column 10 are removed via flow line 104 and combined with hydrogen from flow line 106 into combined feed line 108 and fed to the distillation column reactor 20. In this embodiment the C 's are removed with the C ₃ 'ε in the diεtillation column.

Distillation column reactor 20 is shown to have a stripping εection 26 in the lower half and a rectifying section 28 in the upper half. Two catalyst beds are disposed in the rectifying section. The lower catalyst bed 22 contains the nickel sulfide catalyst in the form of a catalytic distillation εtructure for the mercaptan-diolefin reaction and the upper catalyst bed 24 contains the palladium oxide in the form of a catalytic distillation structure for the selective hydrogenation/isomerization reactionε.

The combined feed εtream in flow line 108 iε fed into the diεtillation column reactor directly below the lower bed. The C ₆ + material is separated from the C ₅ and lighter material in the stripping section 26 with the C ₅ and lighter material boiling up into the first catalyst bed where the diolefins react with substantially all of the mercaptans to form higher boiling sulfides. The sulfides

are diεtilled back down the column into the stripping section where they are removed as bottoms with the C ₆ + heavier material via flow line 110. A portion of the bottoms may be circulated through reboiler 50 via flow line 112 to provide heat balance to the column. The remainder of the bottoms are taken as product via flow line 114.

The substantially sulfur free C ₅ and lighter material is then boiled upward into the upper bed 24 of the rectifying section where the material iε contacted with hydrogen in the preεence of the palladium catalyεt. The remaining diolefinε and acetylenes are selectively hydrogenated to mono-olefins and the mono-olefins are isomerized to equilibrium. The C ₅ and lighter distillate (C ₅ -) , less the mercaptans, diolefins and acetylenes and having an increased percentage of 2-methyl-butene-1 and 2-methyl- butene-2 are removed as overheads via flow line 116 and passed through condenser 30 where the condensible materials are condensed. The liquids are collected in accumulator 40 where the gaseouε materials, including any unreacted hydrogen, are separated and removed via flow line 124. The unreacted hydrogen may be recycled (not shown) if desired. The liquid distillate product is removed via flow line 122. Some of the liquid is recycled to the column 20 as reflux via line 120.

Generally the C ₅ and lighter material will be used aε feed εtock for a etherification unit where the isoamylenes contained therein will be converted to TAME or tertiary amyl ethyl ether (TAEE) . This TAME or TAEE is recombined with the C ₆ bottoms and sent to gasoline blending. If desired the bottoms can be subjected to destructive hydro- deεulfurization to remove the εulfides and other heavier εulfur compounds. In another embodiment a light cracked naphtha aε described is distilled to remove C ₃ 's and H S, with C ₄ •ε and heavier going aε bottomε from the diεtillation column to a firεt diεtillation column reactor containing a dual

bed hydrogenation catalyst, where the C ₅ 's and heavier are taken as bottoms and the C ₄ 's as overheads after contacting the catalysts in the beds with hydrogen. The bottoms may be further treated in a second distillation column reactor with a hydrogenation catalyst and hydrogen to treat the C ₅ portion which is recovered aε the overheads from the second distillation column reactor and the C ₆ and heavier as bottoms. In another embodiment the C ₄ *ε and Cs'ε are taken aε the overheads in the first distillation column reactor and the C ₆ 's and heavier as bottoms.

EXAMPLES In the Examples a three inch diameter column iε loaded with 35 feet of the palladium catalyεt aε distillation structure in the upper portion of the column. Below the first catalyst 13.3 feet of the nickel catalyst was loaded. A stripping section of 50 feet containing Pall rings was left below the lower nickel catalyst bed.

In Example 1 the feed to the reaction diεtillation column waε a C5+ naphtha with the C ₃ /H ₂ S/C removed in the diεtillation column, with the C ₅ 's being taken as overheads after contact with the dual beds. In Example 2 the naphtha cut is C +, with C ₄ /C ₅ being contacted with the dual bedε and being taken as overheads. In both examples the bottoms are C ₆ +. The conditions and resultε are shown in TABLE III below. The chromatographic analysis of the overheads was conducted for undersirables. In both runε the sulfur (mercaptan) reduction was essentially complete and dienes were reduced over 99.8%.

TABLE III

EXAMPLE 1 2 Conditions

Presεure, pεig 130 130

H ₂ partial press, psia 6.00 3.35

Temperature, °F ovhd 226 219.9 top bed 226 226 lower bed 245 245 Flow rates, lbs/hr feed 219.9 219.9 ovhd 56.5 56.5 mdrflx 135.0 135 H ₂ rate, scfh 40.0 40 Feed Analysis total C ₄ 's, wt % 7.67 C ₄ dienes, wt % of C ₄ 'ιs 0.97 butenes, wt % of C ₄ 's 53.85 total C ₅ 's, wt % 26.4 16.4 dienes wt % of C ₅ 's 1.99 1.99 n-pentenes, wt% of Cs'ss 29.92 29.29 isoamyleneε, wt% of C5' •εs 34.65 34.65 Pentene-1:n-pentene % 19.0 19.0 3MB1:IA, % 4.7 4.7

EtSH in C ₅ 'ε (Sulfur) wppm 51 51 MeSH in feed, wppm 36 Overheadε Analyεis Dienes, wt% 0.0046 0.0046 1,3-BD 0.000

EtSH, (Sulfur) , wppm 0 0 MeSH, wppm 0 Pentene-1:n-pentene 5.3 5.3 3MB1.IA 1.2 1.2