WO1994019428A1

WO1994019428A1 - REFORMING USING A Pt LOW-Re CATALYST IN THE LEAD REACTOR

Info

Publication number: WO1994019428A1
Application number: PCT/US1994/002038
Authority: WO
Inventors: Eduardo Mon; William C. Baird
Original assignee: Exxon Research & Engineering Company
Priority date: 1993-02-18
Filing date: 1994-02-18
Publication date: 1994-09-01

Abstract

Catalytic reforming wherein the lead reactor contains a catalyst comprised of Pt and a relatively low level of Re on an inorganic oxide support. The tail reactor contains a platinum rhenium catalyst wherein the rhenium content is at higher levels.

Description

REFORMING USING A Pt LOW-Re CATALYST IN THE LEAD REACTOR

Field of the Invention

The present invention relates to catalytic reformin wherein the lead reactor contains a catalsyt comprised of P and a relatively low level of Re on an inorganic oxid support. The tail reactor contains a platinum-rheniu catalyst containing higer levels of rhenium.

Background of the Invention

Catalytic reforming is a process for improving th octane quality of naphthas or straight run gasolines. Th catalyst is typically multi-functional and contains a meta hydrogenation-dehydrogenation (hydrogen transfer) component or components, composited with a porous, inorganic oxid support, notably alumina. Noble metal catalysts, notably o the platinum type, are currently employed, reforming bein defined as the total effect of the molecular changes, o hydrocarbon reactions, produced by dehydrogenation o cyclohexanes and dehydroisomerization of alkylcyclopentanes t yield aromatics; dehydrogenation of paraffins to yiel olefins; dehydrocyclization of paraffins and olefins to yiel aromatics; isomerization of n-paraffins; isomerization o alkylcycloparaffins to yield cyclohexanes; isomerization o substituted aromatics; and hydrocracking of paraffins whic produces gas, and inevitably coke, the latter being deposite on the catalyst.

Platinum is widely commercially used in th production of reforming catalysts, and platinum-on-alumin catalysts have been commercially employed in refineries fo the last few decades. In the last several years, additiona metallic components have been added to platinum as promoter to further improve the activity or selectivity, or both, o the basic platinum catalyst, e.g. , iridium, rhenium, tin, an the like. Some of the polymetallic catalysts possess superio activity, or selectivity, or both, as contrasted with othe catalysts. Platinum-rhenium catalysts by way of exampl possess admirable selectivity as contrasted with platinu catalysts, selectivity being defined as the ability of th catalyst to produce high yields of C5+ liquid products wit concurrent low production of normally gaseous hydrocarbons, i.e., methane and other gaseous hydrocarbons, and coke. Iridium-promoted catalysts, e.g., platinum-iridium, an platinum-iridium-tin (U.S. 4,436,612) catalysts, on the othe hand, are known for their high activity, as contrasted e.g., with platinum and platinum-rhenium catalysts, activity bein defined as the relative ability of a catalyst to convert given volume of naphtha per volume of catalyst to high octan reformate.

In a reforming operation, one or a series o reactors, or a series of reaction zones, are employed. Typically, a series of reactors is employed, e.g., three o four reactors, these constituting the heart of the reformin unit. Each reforming reactor is generally provided with fixed bed, or beds, of the catalyst which receive downflo feed, and each is provided with a preheater or interstag heater, because the reactions which take place ar endothermic. A naphtha feed, with hydrogen, or recycl hydrogen gas, is passed through a preheat furnace and reacto and then in sequence through subsequent interstage heaters an reactors of the series. The product from the last reactor i separated into a liquid fraction, and a vaporous effluent. The former is recovered as a C5+ liquid product. The latte is a gas rich in hydrogen, and usually contains small amount of normally gaseous hydrocarbons, from which hydrogen i separated and recycled to the process to minimize cok production. The sum-total of the reforming reactions, supra occurs as a continuum between the first and last reactor the series, i.e., as the feed enters and passes over the firs fixed catalyst bed of the first reactor and exits from th last fixed catalyst bed of the last reactor of the series. The reactions which predominate between the several reacto differ dependent principally upon the nature of the feed, an the temperature employed within the individual reactors. I the initial or lead reactor, which is maintained at a rela tively low temperature, it is believed that the primar reaction involves the dehydrogenation of naphthenes t produce aromatics. The isomerization of naphthenes, notabl C5 and C6 naphthenes, also occurs to a considerable extent. Most of the other reforming reactions also occur, but only t a lesser, or smaller extent. There is relatively littl hydrocracking, and very little olefin or paraffi dehydrocyclization occurring in the first reactor. Withi the intermediate reactor zone(s), or reactor(s) , th temperature is maintained somewhat higher than in the first or lead reactor of the series, and it is believed that th primary reactions in the intermediate reactor, or reactors involve the isomerization of naphthenes and paraffins. Where e.g. , there are two reactors disposed between the first an last reactor of the series, it is believed that the principa reaction involves the isomerization of naphthenes, norma paraffins and isoparaffins. Some dehydrogenation o naphthenes may, and usually does occur, at least within th first of the intermediate reactors. There is usually som hydrocracking, at least more than in the lead reactor of th series, and there is more olefin and paraffin dehydrocycliz ation. The third reactor of the series, or secon intermediate reactor, is generally operated at a somewha higher temperature than the second reactor of the series. I is believed that the naphthene and paraffin isomerizatio reactions continue as the primary reaction in this reactor but there is very little naphthene dehydrogenation. There i a further increase in paraffin dehydrocyclization, and mor hydrocracking. In the final reaction zone, or final reactor, which is operated at the highest temperature of the series, i is believed that paraffin dehydrocyclization, particularly th dehydrocyclization of the short chain, notably C6 and c paraffins, is the primary reaction. The isomerizatio reactions continue, and there is more hydrocracking in thi reactor than in any one of the other reactors of the series.

The activity of the catalyst gradually declines du to the build-up of coke. Coke formation is believed to resul from the deposition of coke precursors such as anthracene, coronene, ovalene, and other condensed ring aromatic molecule on the catalyst, these polymerizing to form coke. Durin operation, the temperature of the the process is graduall raised to compensate for the activity loss caused by the cok deposition. Eventually, however, economics dictate th necessity of reactivating the catalyst. Consequently, in all processes of this type the catalyst must necessarily b periodically regenerated by burning of the coke at controlle conditions.

Improvements have been made in such processes, an catalysts, to reduce capital investment or improve C5+ liqui yields while improving the octane quality of naphthas and straight run gasolines. New catalysts have been developed, old catalysts have been modified, and process conditions have been altered in attempts to optimize the catalytic contribution of each charge of catalyst relative to a selected performance objective. Nonetheless, while any good commercial reforming catalyst must possess good activity, activity maintenance and selectivity to some degree, no catalyst can possess even one, uchless all of these properties to the ultimate degree. Nonetheless, while catalysts with high activity are very desirable, there still remains a need, and indeed a high demand, for increased selectivity; and even relatively small increases in C5+ liquid yield can represen large cr edits in commercial reforming operations. Further since the advent of blending oxygenates into refinery moga pools, many catalytic reforming units will be driven toward lower reformate octanes. This will result in lower hydroge yields. Consequently, a need exits for catalyst which ar more selective for hydrogen make.

Although a large number of various reformin catalysts and processing schemes have been developed over th years, there is still a need in the art for more effecient an selective operation of commercial reforming units which tak advantage of the properties of a particular catalyst.

Summary of the Invention

In accordance with the present invention, there i provided a process for reforming a naphtha feedstream t obtain an improved C₅+ liquid yield, which process comprise conducting the the reforming in a series of reactors wherein (a) the lead reactor contains a catalyst comprised of abou 0.1 to 1 wt.% Pt and about 0.01 to 0.1 wt.% Re, on a inorganic oxide support; and

(b) the tail reactor contains a catalyst compris of about 0.1 to 1 wt.% Pt, and from about 0.1 percent to abo 1.0 percent rhenium, based on the total weight of the cataly (dry basis) .

In a preferred embodiment of the present inventi the catalyst of the lead reactor contains from about 0.2 0.7 wt.% Pt and about 0.02 to 0.07 wt.% Re.

Detailed Description of the Invention

As previously stated, the present invention relat to reforming naphtha feedstocks boiling in the gasoline rang Non-limiting examples of such feedstocks include a virg naphtha, cracked naphtha, a naphtha from a coal liquefactio process, a Fischer-Tropsch naphtha, or the like. Typica feeds are those hydrocarbons containing from about 5 to abou 12 carbon atoms, or more preferably from about 6 to about carbon atoms. Naphthas, or petroleum fractions boiling withi the range of from about 25°C. to about 230°C. , and preferabl from about 50°C. to about 190°C. , contain hydrocarbons o carbon numbers within these ranges. Typical fractions thu usually contain from about 15 to about 80 vol. % paraffins both normal and branched, which fall in the range of about C to C₁₂, from about 10 to 80 vol. % of naphthenes fallin within the range of from about CQ to C₁₂, and from 5 throug 20 vol. % of the desirable aromatics falling within the rang of from about C₆ to C*-^-

The reforming is conducted in a reforming proces unit comprised of a plurality of serially connected reactors For purposes of the present invention, it is important tha the lead, or first, reactor contain a catalyst comprised o about 0.1 to 1 wt.% of Pt, preferably from about 0.2 to 0. wt.% Pt; and about 0.01 to 0.1 wt.% Re, preferably from abou 0.02 to 0.07 wt.% Re, on an inorganic oxide support. Th weight percents are based on the total weight of the catalys (dry basis) .

Reforming reactions in the tail reactor ar typically paraffin dehydrocyclization and hydrocracking an for purposes of the present invention are conducted in th presence of a catalyst comprised of about 0.1 to 1 wt.% Pt preferably from about 0.2 to 0.7 wt.% Pt; and about 0.1 to wt.% Re, preferably from about 0.2 to 0.7 wt.% Re, also base on the total weight of the catalyst (dry basis) . The metal of this catalyst will be substantially uniformly disperse throughout the support. The catalyst used in the present invention will preferably also contain halogen, preferably chlorine, in concentration ranging from about 0.1 percent to about 3 percent, preferably from about 0.8 to about 1.5 percent, based on the total weight of the catalyst. It is also preferred that the catalyst be sulfided, e.g., by contact with a hydrogen sulfide-containing gas, and contains from about 0.01 percent to about 0.2 percent, more preferably from about 0.05 percent to about 0.15 percent sulfur, based on the total weight of the catalyst. The metal components, in the amounts stated, are uniformly dispersed throughout an inorganic oxide support, preferably an alumina support and more preferably a gamma alumina support.

The relative loadings of each catalyst should be such that they are sensitive to feed type and process conditions. The distribution of the catalyst types between lead and tail reactors may be varied as desired. In general, the catalyst in the tail reactors will a-fccount for about 20 to 90 wt.%, preferably from about 30 to 80 wt.%, and more preferably from about 50 to 70 wt.%, based on the total amount of catalyst charged to the reforming unit.

Practice of the present invention results in the suppression of excessive dealkylation reactions with simultaneous increase in dehydrocyclization reactions to increase C5+ liquid yields, with only a modest activity debit vis-a-vis the use of a catalyst in the tail reactor which is otherwise similar but does not contain the tin, or contains tin in greater or lesser amounts than that prescribed for the tail reactor catalyst of this invention. In addition to the increased C5+ liquid yields, temperature runaway rate during process upsets is tempered, and reduced; the amount of benzene produced in the reformate at similar octane levels is reduced, generally as much as about 10 percent to about 15 percent, based on the volume of the C5+ liquids, and there is lowe production of fuel gas, a product of relatively low value.

Practice of the present invention results in th suppression of execessive dealkylation reactions wit simultaneous increase in dehydrocyclization reactions t increase C₅+ liquid yields. This is accomplished with only modest activity debit vis-a-vis the use of a catalyst in th tail reactor which is otherwise similar but does not contai the low levels of Re. In addition to the increased C₅+ liqui and hydrogen yields, temperature runaway rate during proces upsets is tempered, and reduced. The amount of benzen produced in the reformate at similar octane levels is reduced, generally as much as about 10 percent to about 15 percent, based on the volume of the C5+ liquids. There is also lowe production of fuel gas, a product of relatively low value.

The catalyst employed in accordance with thi invention is necessarily constituted of composite particle which contain, besides a support material, th hydrogenation-dehydrogenation components, a halide componen and, preferably, the catalyst is sulfided. The suppor material is constituted of a porous, refractory inorgani oxide, particularly alumina. The support can contain, e.g., one or more alumina, bentonite, clay, diatomaceous earth, zeolite, silica, activated carbon, magnesia, zirconia, thoria, and the like; though the most preferred support is alumina t which, if desired, can be added a suitable amount of othe refractory carrier materials such as silica, zirconia, magnesia, titania, etc., usually in a range of about 1 to 2 percent, based on the weight of the support. A preferre support for the practice of the present invention is on having a surface area of more than 50 m²/g, preferably fro about 100 to about 300 m²/g, a bulk density of about 0.3 t 1.0 g/ml, preferably about 0.4 to 0.8 g/ml, an average por volume of about 0.2 to 1.1 ml/g, preferably about 0.3 to 0.8 ml/g, and an average pore diameter of about 30 to 300 Angstro units.

The metal hydrogenation-dehydrogenation component can be uniformly dispersed throughout the porous inorgani oxide support by various techniques known to the art such a ion- exchange, coprecipitation with the alumina in the sol o gel form, and the like. For example, the catalyst composit can be formed by adding together suitable reagents such as salt of tin, and ammonium hydroxide or carbonate, and a sal of aluminum such as aluminum chloride or aluminum sulfate t form aluminum hydroxide. The aluminum hydroxide containin the tin salt can then be heated, dried, formed into pellets o extruded, and then calcined in air or nitrogen up to 1000°F. The other metal components can then be added. Suitably, th metal components can be added to the catalyst by impregnation typically via an "incipient wetness" technique which require a minimum of solution so that the total solution is absorbed initially or after some evaporation.

To enhance catalyst performance in reformin operations, it is also required to add a halogen component t the catalysts, fluorine and chlorine being preferred haloge components. The halogen is contained on the catalyst withi the range of 0.1 to 3 percent, preferably within the range o about 0.8 to about 1.5 percent, based on the weight of th catalyst. When using chlorine as the halogen component, it i added to the catalyst within the range of about 0.2 to percent, preferably within the range of about 0.8 to 1. percent, based on the weight of the catalyst. The introduc tion of halogen into the catalyst can be carried out by an method at any time. It can be added to the catalyst durin catalyst preparation, for example, prior to, following o simultaneously with the incorporation of a meta hydrogenation-dehydrogenation component, or components. It ca also be introduced by contacting a carrier material in a vapo phase or liquid phase with a halogen compound such as hydroge fluoride, hydrogen chloride, ammonium chloride, or the like.

The catalyst is dried by heating at a temperatur above about 25°C. , preferably between about 65°C. and 175°C. , in the presence of nitrogen or oxygen, or both, in an ai stream or under vacuum. The catalyst is calcined at temperature between about 200°C. to 450°C. , either in th presence of oxygen in an air stream or in the presence of a inert gas such as nitrogen.

Sulfur is a highly preferred component of th catalysts, the sulfur content of the catalyst generall ranging to about 0.2 percent, preferably from about 0.05 percent to about 0.15 percent, based on the weight of th catalyst (dry basis) . The sulfur can be added to the catalys by conventional methods, suitably by breakthrough sulfiding o a bed of the catalyst with a sulfur-containing gaseous stream, e.g., hydrogen sulfide in hydrogen, performed at temperature ranging from about 350°F. to about 1050°F. , and at pressure ranging from about 1 to about 40 atmospheres for the tim necessary to achieve breakthrough, or the desired sulfu level.

The reforming runs are initiated by adjusting th hydrogen and feed rates, and the temperature (Equivalen Isothermal Temperature) and pressure to operating conditions. The run is continued at optimum reforming conditions b adjustment of the major process variables, within the range described below: LEAD REACTOR CONDITIONS

Major Operating Typical Process Preferred Process Variables Conditions Conditions

Pressure, psig 100-700 150-500

Reactor Temp. , °C. 370-540 425-510 Recycle Gas Rate, SCF/B 2000-10,000 2000-6000 Feed Rate, W/Hr/W 1-20 2-10

TAIL REACTOR CONDITIONS

Pressure, psig 100-700 150-500

Reactor Temp. , °C. 425-565 450-525 Recycle Gas Rate, SCF/B 2000-10,000 2000-6000 Feed Rate, W/Hr/W 1-10 2-8

The invention will be more fully understood by reference to the following comparative data illustrating its more salient features. All parts are given in terms of weight except as otherwise specified.

In conducting these tests, an n-heptane feed was used in certain instances. In others a full range naphtha was employed.

Inspections on the full range Arab Light Naphtha feed employed in making certain of the tests are given below. Property Arab Light Naphtha

Gravity at 15°C

API⁰ 59.4

Specific 0.7412 Octane, RON Clear 38 Molecular Weight 111.3 Sulfur, wppm 0.3

Distillation D-86,

IBP 89, 5% 102,

10% 105,

50% 125,

90% 153,

95% 160,

FBP 171, Composition, Wt. % Total Paraffins 65.1 Total Naphthenes 19.3 Total Aromatics 15.6

Example 1

A conventional 0.3 wt.% Pt-0.3 wt.% Re catalyst wa calcined in air at 500°C, reduced in hdyrogen at 500°C for 1 hr. , and sulfided to breakthrough at 500°C with a hydroge with a hydrogen/hydrogen sulfide blend. The catalyst wa tested in heptane reforming, with the results appear in Tabl I below.

Example 2

A 0.3 wt.% Pt, 0.05 wt.% Re catalyst was prepared b the following procedure. Alumina extrudates were suspended i water and carbon dioxide was bubbled through the mixture fo 30 minutes. Solutions of chloroplatinic acid, perrhenic acid, and hydrochloric acid were added in the appropriat quantities, and the mixture was treated with carbon dioxid for 4 hours. The extrudates were dried, and the catalyst wa calcined in air for 3 hours, reduced in flowing hydrogen fo 17 hours, and sulfided with a hydrogen-hydrogen sulfide blend, all at 500°C. This catalyst was tested in heptane reforming and the results are shown in Table I below.

Table I n-Heptane. 500°C. 100 psiσ. 10 W/H/W. H₂/Oil-6

Catalyst

Yield, wt.% on feed 0.3Pt-0.3Rβ 0.3Pt-0.05Re l 1.4 1.1 i-c₄ 3.8 2.7 n-C₄ 5.6 3.7 c₅+ 78.9 85.2

Toluene 28.5 30.1

Conversion 65.2 57.3

Toluene Rate 2.9 3.1

Toluene Selectivity 43.7 52.5

The above data show that the Pt-low concentration Re catalyst used in the lead reactor in the present invention is more selective than the conventional Pt-Re catalyst in terms of higher C5+ liquid yield and toluene selectivity. The Pt- low concentration Re catalyst and the conventional Pt-Re catalyst are substantially at parity in terms of activity. The selectivity credits for the low Re catalyst used in the lead reactor are evident when the catalysts are tested on a full range naphtha at conditions simulating those in a commercial lead reactor. These data are presented in Table II below.

Table II

Lead Reactor Reforming of Light Arab Paraffinic Naphtha at 500°C. 350 psiσ. 4500 SCF/B. 1.4 W/H/W

Catalyst 0.3Pt-0.3Re 0.3Pt-0.05Re

Octane 96 96

C5+ LV% @ 100 ROΛ/ 62 70 The results demonstrate that at lead reacto conditions, the activities of the Pt-Re catalysts ar susbstantially at parity. However, the selectivity advantag offerred by the Pt-low Re catalyst provides a substantia yield credit, and for this reason the Pt-low Re catalyst show unexpected results over the conventional Pt-Re catalyst whe used in the lead reactor.

Example 3

The 0.3 wt/% Pt, 0.05 wt.% Re on alumina catalyst o Example 2 was staged with a 0.3 wt.% Pt, 0.3 wt.% Re o alumina catalyst is a single isothermal reactor. A naphth feedstock was introduced into the reactor so that the low R catalyst represented the first stage and the conventional 0. wt.% Re catalyst the second stage. The feedstock had boiling range from about 90°C to about 150°C and was comprise of about 55.3 wt.% paraffins, 28.5 wt.% naphthenes, and 16. wt.% aromatics. Table III below gives the test conditions an the resulting hydrogen and C5+ yields, and relative activity.

Table III

316 psiσ. 1.2 WHW. H₂/Oil = 1.9

Catalyst 0.3 Pt-0.3 Re 0.3 Pt-0.05 Re/0.3 Pt-0.3 R

Yields § 100 RONC

H₂ Wt.% 1.6 1.77

C₅+, LV% 70.9 73.0

Relative Activity 100 94

These results again demonstrate that at lead reacto conditions, the activities of the Pt-Re catalysts ar substantially at parity. However, the selectivity advantag offered by the Pt-low Re catalyst provides a substantial C5 liquid and hydrogen yield credit, and for this reason the Pt low Re catalyst shows unexpected results over the conventiona Pt-Re catalyst when used in the lead reactor.

Claims

CLAIMS :

1. A process for reforming a naphtha feedstream to obtain an improved C5+ liquid and hydrogen yield, which process comprises conducting the the reforming in a series of reactors wherein:

(a) the lead reactor contains a catalyst comprised of about 0.1 to 1 wt.% Pt and about 0.01 to 0.1 wt.% Re on an inorganic oxide support; and

(b) the tail reactor contains a catalyst comprised of about 0.1 to 1 wt.% Pt, from about 0.1 percent to about 1 wt.% Re based on the total weight of the catalyst.

2. The process of claim 1 wherien the catalyst of the lead reactor contains from about 0.2 wt. % to 0.7 wt.% Pt, and from about 0.02 wt.% to 0.07 wt.% Re.

3. The process of Claim 1 wherein the catalyst of the tail reactor contains from about 0.2 wt.% to about 0.7 wt.% Pt and from about 0.2 percent to about 0.7 wt.% Re.

4. The process of Claim 1 wherein the catalyst contains from about 0.1 percent to about 3.0 percent halogen.

5. The process of Claim 1 wherein the catalyst contains from about 0.01 percent to about 0.2 percent sulfur.

6. The process of Claim 1 wherein the inorganic oxide support component of the catalyst is alumina.

7. The process of Claim 1 wherein the reforming conditions employed in the tail reactor of the series are defined as follows:

Pressure, psig about 100 to 700

Reactor Temperature, about 425 to 565 Gas Rate, SCF/B about 2000 to 10,000 Feed Rate, W/Hr/W about 1 to 10.

8. The process of Claim 7 wherein the reforming conditions employed in the tail reactor of the series are defined as follows:

Pressure, psig about 150 to 500

Reactor Temperature, °C about 450 to 525 Gas Rate, SCF/B about 2000 to 6000 Feed Rate, W/Hr/W about 2 to 8.

9. The process of Claim 1 wherein the reforming conditions employed in the lead reactors of the series are defined as follows:

Pressure, psig about 100 to 700

Reactor Temperature, about 370 to 540 Gas Rate, SCF/B about 2000 to 10,000 Feed Rate, W/Hr/W about 1 to 20.

10. The process of Claim 9 wherein the reforming conditions employed in the lead reactors of the series are defined as follows:

Pressure, psig about 150 to 500

Reactor Temperature, °C about 425 to 510

Gas Rate, SCF/B about 2000 to 6000

Feed Rate, W/Hr/W about 2 to 10,