Lloyd L. Handbook of Industrial Catalysts

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238 Chapter 6

6.6. CATALYTIC REFORMING

Catalytic reforming is a key reﬁnery process. It improves the octane rating of

virgin naphthas and light distillates so they can be used in gasoline formulations.

The process has also become an important source of aromatics for use in petro-

chemical production.

Low-octane, sulfur-and nitrogen-free parafﬁns and naphthas in the boiling

range 80–160

C are converted to higher octane iso-parafﬁns and aromatic com-

pounds by a sequence of dehydrogenation, dehydrocyclisation and isomerization

reactions. Hydrocracking reactions are less important, but can also convert the

higher-molecular-weight paraffins to more useful products, but with a reduction

in the yield of C

products.

When gasoline is the main product required, it is usual to remove the C

–C

straight-run naphtha cut before reforming the 100–160

C fraction. This is be-

cause separate isomerization of the C

–C

paraffins (Section 6.10) gives an

overall improvement in octane number. If aromatics are required, a lower-

boiling fraction, in the range 80–120

C, is reformed. The heavy naphthas are not

reformed because coke forms more readily and quickly leads to deactivation of

the catalyst.

The ﬁrst naphtha-reforming catalysts were used during World War II in a

process known as hydroforming and were based on molybdenum or chromium

oxides supported on alumina.

At the time, the new process provided high-

octane aviation fuel and the toluene needed for the production of explosives.

The decreased demand for these products after the war meant that the seven op-

erating plants were idle by 1952. In any case the preferred catalyst (9%

MoO

/alumina) was neither active enough nor selective enough to give a really

economic performance and needed frequent regeneration. Vladimir Haensel,

who invented the platinum reforming catalyst,

recalled that during and after the

wartime period, UOP was investigating a hydrocracking process that used desul-

furized kerosene as feed. The work involved an analytical procedure for six-

membered rings, based on the catalytic dehydrogenation of naphthenes with 3%

platinum supported on carbon. Haensel tells how this led to the treatment of

virgin naphthas containing sulfur with a similar platinum catalyst supported on

silica. Operation at up to 450

C and 500 psig, with 5 mol of hydrogen recycle

per mole of feed, showed not only that the virgin straight-run naphthas were

reformed to give a high-octane gasoline product but that sulfur poisoning was

avoided. At the same time, however, the silica support and others such as sili-

ca/alumina promoted hydrocracking reactions. Further investigations produced

an economic platinum/alumina catalyst with a platinum content in the range 0.3–

0.7%.

The catalyst proved to be particularly effective when the alumina was

prepared from aluminum chloride rather than the nitrate. This led to the treat-

ment of supports with chloride or fluoride to improve acidity—a practice that

has continued ever since. The new UOP process was announced at the Western

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240 Chapter 6

6.6.1. Naphtha Reforming Reactions

The principal reforming reactions are shown in Table 6.14. The isomerization of

n-paraffins proceeds via a dehydrogenation step on the platinum component of

the catalyst to an n-oleﬁn intermediate. The n-oleﬁn then migrates to the acidic

alumina site and isomerizes to the iso-oleﬁn. Finally the iso-olefin is hydrogen-

ated to the isoparaffin by the platinum. This mechanism requires good disper-

sion of platinum on the highly acidic, chloride-treated alumina support.

Similar mechanisms apply to the isomerization and dehydrocyclization of n-

hexane to methyl cyclopentane and then benzene, as shown in Table 6.15.

6.6.1.1. Reformer Operation

The combined sequence reforming reaction is endothermic, and there is a

signiﬁcant temperature decrease across the catalyst bed. Heat has to be supplied

continuously to maintain operation, so catalyst is loaded into a number of beds

or reactors with interbed heating of the reaction mixture. The process is compli-

cated because of the different reactions and temperature changes taking place in

each bed. Variable heat balances were compensated in early plants by loading

increasing volumes of catalyst in each bed from the top to the bottom of each

reactor. An alternative arrangement was to load several reactors with equal vol-

umes of catalyst while having a similar reactor in reserve to replace any operat-

ing reactor needing regeneration.

Three different reactor systems have evolved since the process was ﬁrst in-

troduced in 1952:

Table 6.14. Naphtha Reforming Reactions.

Reaction Heat balance

Dehydrogenation:

Naphthenes to aromatics

(e.g., methylcyclohexane to toluene)

Extremely endothermic

Dehydroisomerization:

(e.g., methyl cyclopentane to benzene)

Endothermic

Isomerization:

Normal paraffins to branched parafﬁns

Mildly exothermic

Dehydrocyclization:

Normal paraffins to aromatics

Endothermic

Hydrocracking:

Normal heptane to C

and C

paraffins

Exothermic

TABLE 6.15. Catalytic Rreforming of n-Hexane and iso-Hexanes.

↑

Hydrogenation and

dehydrogenation on

platinum catalyst

↓

n-hexane

iso-hexane



cyclohexane

methylcyclopentane

n-hexene

iso-hexene



(n-hexene carbenium ion)

(iso-hexyl carbenium

ion)

cyclohexene

methylcyclopentene



cyclohexadiene

methylcyclopentadiene



benzene

←Isomerization on acidic alumina→

Note: From Mills, Heinemann, Milliken and Oblad, Ind. and Eng. Chemistry, 1953.

Refinery Catalysts 22 41

242 Chapter 6

• The semi-regenerative process operates continuously over long periods of

up to 9–12 months. There are usually three reactors in series that contain

about 15–20%, 30–35%, and 50–60% respectively of the catalyst’s

charge, depending on the feed used. Operation is in the range 10–25 atm

for different feeds and reforming severity. When the catalyst activity has

decreased, resulting in lower liquid yields and more hydrocracking activi-

ty, the catalyst is regenerated in situ, using air at about 8 atm. The catalyst

can usually be regenerated up to ten times before the performance be-

comes uneconomic and it has to be replaced. Reaction conditions are

summarized in Table 6.16.

• Cyclic processes generally operate with four reactors although as many as

six are used if frequent regeneration is necessary. All reactors contain

equal volumes of catalyst and operate at about 480

C and 20 atm, with a

temperature decrease of 25

–30

C across each bed. Low hydrogen/feed

ratios can be used. One reactor is kept in reserve so that it can be

switched to replace any on-line reactor requiring regeneration, so giving

continuous operation. Cyclic processes operate at low pressure to give

high yields of a high-octane product, but with relatively high coke lay-

down. The interval between switching reactors to regenerate the catalyst

can be rather short, for example, after processing only about eight barrels

of feed per pound of catalyst used, with catalyst regeneration either daily

or perhaps after only a few days. This complicates the process engineer-

ing because alternating hydrogen and oxygen streams have to be provid-

ed. Catalyst life is also relatively short as a result of the frequent regener-

ationneeded and can be less than about 200 barrels of feed per pound of

catalyst.

• Contin uous catalyst regeneration (CCR) processes were introduced in

1971 to avoid the shut-down intervals necessary in semi-regenerative re-

formers. The reformers have four beds, either stacked vertically or placed

side by side horizontally, so that the catalyst can be moved continuously

from one catalyst bed to another and ultimately, into the regenerator (Fig.

6.5). After regeneration, the catalyst is reduced before being returned to

the ﬁrst reactor to begi

n the sequence again. The average residence time

of a catalyst in a CCR unit has decreased since the process was ﬁrst intro-

duced. The ﬁrst design in 1971 recycled the catalyst charge about 12

times a year. By 1980 the recycle rate was 50 times a year, but as operat-

ing conditions have become even more severe, the rate can now be as

high as 150 times a year.

Despite this, the catalysts used have lasted for

more than seven years. Operating conditions are shown in Table 6.17.

Refinery Catalysts 243

TABLE 6.16. Semi-Regenerative Catalytic Reforming.

Catalyst

(vol %)

Temperature (

RON

out

Carbon end of

cycle (%)

Inlet outlet

Reactions

Reactor 1 15–20 500 430 65 3–10

Dehydrogenation

Dehydroisomerization

Reactor 2 30–35 500 470 80 5–15

Dehydrogenation

Dehydroisomerization

Some hydrocracking

Dehydrocyclization

Reactor 3 50 500 495 95–100 8–20

Hydrocracking

Dehydrocyclization

Conditions

Feed (m

−1

) 85–90

Boiling range (

C) ~100–160

/HC 4–5

LHSV (h

−1

) 1–3

Pressure (atm) 10–25

Recycle gas (% H

) 80

Yield reformate (wt%) 70–80

Hydrogen (wt%) 2–3

Feed/products Feed Products

Parafﬁns (%) 45–55 30–50

Naphthenes (%)

30–40 5–10

Aromatics (%) 5–10 45–60

244

Figure

ermis

Through

Pressure

Hydrog

Regener

Liquid s

Catalyst

RON

Temper

Yield H

(vol

Aromati

Chapter 6

6.5. Modern c

ion of UOP).

ut (bbl/day)

(atm)

n/hydrocarbon

tion (lb h

-1

)

tions per yea

ace velocity (h

-1

)

charge (tonnes)

ture (

(wt%)

)

s (vol%)

talytic reformi

LE 6.17. Cont

g plant with c

nuous Catalyti

ntinuous cataly

Regeneration.

971

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200

1.6

450

t regeneration

(

1990

25,000

3–4

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4500

150

1.6

110

102

510

(

Refinery Catalysts 245

The CCR process has been very successful and is now used extensively

throughout the world. The main advantage is that the coke content of catalyst

before regeneration is as low as 5–6%, despite the severe operating conditions at

hydrogen pressure as low as 3 atm. Yields of high-octane product and hydrogen

were significantly increased. Old catalyst can also be withdrawn from the reac-

tor and replaced with fresh catalyst when activity declines. High yields are fa-

vored by low pressure and low hydrogen circulation, but more coke is formed.

Despite the higher capital cost of CCR reformers the higher yields and more

flexible operation make the process very popular.

By the late 1990s more than 900 reformers were operating throughout the

world.

The semi-regenerative process was being used in 60–70% of them, with

continuous regeneration processes being used in 15–20%.

Cyclic processes

were only used in about 10–15% of the reformers. Since 1990, however, almost

all of the new reformers being built or designed use the continuous regeneration

process. As larger-capacity reformers have been built, axial ﬂow through the

catalyst beds has been replaced by radial ﬂow because this configuration allows

for better gas distribution and reduced pressure drop.

Catalyst operation in the catalytic reforming process is influenced by feed

composition, hydrogen ratio, and operating conditions. The beneﬁts obtained by

any changes in design always have to be balanced with certain disadvantages.

For example, although straight-run naphthas boiling at up to about 200

C can be

used in catalytic reforming processes, the use of lower-boiling naphthas gives

better results. More coke forms on the catalyst as high-molecular-weight hydro-

carbons are reformed, and it is estimated that above 190

C every 13

C increase

in the boiling range of the naphtha can reduce the period before regeneration by

35%.

Sulfur impurities in naphtha poison platinum hydrogenation activity and ni-

trogen impurities neutralize acid sites. Both poisons have to be carefully re-

moved before the naphtha can be used. Despite this need to remove sulfur com-

pounds from reformer feeds, it is also necessary to selectively poison the active

platinum in new or regenerated catalysts by adding about 0.25% of sulfur to the

catalyst in a sulﬁding procedure using 100–200 ppm H

S in the feed. This limits

the extent to which the exothermic hydrogenolysis reactions that deposit coke

and lead to sintered catalyst can occur. Without this sulfur conditioning, the

yield of product can be reduced by 1% and the run length reduced by up to 15–

20%. Unlike chloride, sulfur should not be added continuously during operation

and the normal sulfur speciﬁcation of naphtha must be applied.

Since the ﬁrst catalytic reformers were used in the 1950s the hydrogen/ hy-

drocarbon mole ratio and the reformer operating pressure have both been gradu-

ally decreased. These developments resulted from improvements in the alumina

support, the use of bimetallic catalysts, and ﬁnally the introduction of the low-

pressure, continuous catalyst regeneration processes. The trends in operation are

shown in Table 6.18.

246 Chapter 6

TABLE 6.18. Trends in Reformer Severity.

These changes have resulted in improved operation, not only by reducing

gas compression costs, but also by increasing gasoline yields and hydrogen pro-

duction from the dehydrogenation reactions, in combination with a lower degree

of the hydrocracking reaction. Unfortunately, some hydrogen is lost because it

also reacts (beneﬁcially) with coke precursors, so removing them and decreasing

the rate of coke deposition. The use of low hydrogen pressure slows down this

reaction, and ultimately leads to more coke deposition and the need for more

frequent regeneration. It is necessary, therefore, to balance the advantage of

higher yields against increased catalyst deactivation in processes that allow con-

tinuous catalyst regeneration. Modern catalytic reformers operate at liquid hour-

ly space velocity in the range 1.0–2.0 h

-1

. The lower limit is a compromise be-

tween an increase in undesirable hydrocracking and a decrease in beneﬁcial de-

hydrocyclization reactions. These conditions have no effect on either dehydro-

genation or isomerization reactions, which still reach equilibrium. The inlet

temperature to the catalyst bed must be increased as the catalyst is either deac-

tivated or if a higher octane product is required. Operating temperatures also

depends on the feed composition, hydrogen ratio, and the need to control the

temperature decrease in all beds. Typical temperatures are shown in Tables 6.16

and 6.17.

An essential requirement of reformer operation is the addition of chlorine

compounds to the feed to maintain the necessary chlorine content and hence the

acidity of the alumina support. This requires a careful monitoring of the mass

balance of chloride entering and leaving the catalyst beds, to maintain the chlo-

ride content of the catalyst at about 1.0 wt%. A strict control of water in the feed

is also required to minimize chloride loss.

6.6.1.2. Coke Formation

Hydrogenolysis reactions cause the formation of coke precursors on the active

platinum surface, which can then diffuse to adjacent acid sites on the alumina

support. A series of polymerization and dehydrogenation reactions form coke,

which eventually covers the catalyst surface and reduces activity. Early experi-

ence showed that traces of sulfur in the catalyst during start-up reduced coking

by inhibiting the excessive hydrogenolysis reaction and the interval before re-

generation became necessary could be increased. Acid catalyzed cracking reac-

Year H

/hydrocarbon Space velocity (h

−1

) Pressure (atm)

1950s

1960–1970

1970–1990

8–10

5–8

2–5

2–3

1–2

25–35

15–25

3–15

Refinery Catalysts 247

tions were not affected, however, and some coke was still formed. No sulfur

addition is required during normal operation because it shortens the cycle time.

The interval between regenerations can be extended by increasing hydrogen

recycle and raising operating pressure from 15 to 35 atm. This improves the rate

at which coke precursors are hydrogenated on the platinum surface, but these

changes decrease yields and octane ratings (see Section 6.6.1.1).

6.6.2. Reforming Catalysts

Platinum was not only the most active and selective Group-VIII metal for the

naphtha reforming process, but was also more readily available and relatively

less expensive than iridium or rhodium.

Both γ- and η-aluminas were suitable as the acidic support. Although η-

alumina is more acidic and has a higher isomerization activity, it is more expen-

sive to produce and not much better than a stable γ-alumina to which chloride

has been added during its preparation. The composition of typical reforming

catalysts is shown in Table 6.19. Fluorided-alumina and silica/alumina supports

were both used in some early reformers but were not really successful because

of excess acidity, which promoted cracking reactions.

The success of platforming soon led to the rapid introduction of competitive

processes as demand for higher-octane gasoline for high-compression automo-

bile engines increased. Seven process licensors are still offering reforming pro-

cesses differing only in the physical appearance of the unit and the catalysts

used.

More than 50 different catalyst types are currently being supplied by at

least ten catalyst producers. As usual with licensed processes, full details of

catalyst compositions and operation are not described in detail.

Table 6.19. Typical Reforming Catalysts.

Catalyst Properties

Platinum Early type, Pt 0.65% Later type, Pt 0.35%

Platinum/rhenium Balanced , Pt 0.3% / Re 0.3% Skewed , Pt 0.3%/Re 0.6%

Platinum/tin CCR units, Pt 0.3% / Sn 0.6% No change

Platinum/iridium Exxon units, Pt 0.3% / Ir 0.3% No change

Support

-Al

Surface area (m

−1

) 200

Pore volume (ml g

−1

) 0.7

Pore diameter (nm) 6–7

Reactor bulk density (kg liter

−1

) 0.5–0.6 (normal)

0.7–0.8 (dense packed)

Chloride (wt%) 0.9–1.2