Lloyd L. Handbook of Industrial Catalysts

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358 Chapter 9

of the hydrodesulfurization catalysts and absorbents although larger volumes of

absorbent must be included in the plant design to achieve a convenient operating

cycle before replacement.

9.3.1 Activated Carbon

High-surface area, activated carbon has an affinity for some organo-sulfur com-

pounds and can be used as an absorbent for sulfur in chemical processes. For

example, some of the early, low pressure ammonia plants operating in North

America used two parallel beds of activated carbon to adsorb any sulfur impuri-

ties from natural gas. One bed operated while the second bed was being regener-

ated with steam to allow continuous ammonia production. However, mercaptans

and carbonyl sulfide are not strongly bound to activated carbon and can be dis-

placed by higher molecular weight hydrocarbons in the gas stream. To improve

adsorption efficiency the carbon was impregnated with iron or copper oxides

which combined with sulfur but could be easily regenerated. Even so, sulfur

retention was not reliable enough for the activated carbons to be used in modern

high-pressure ammonia plants. The high cost of regular regeneration, plus sulfur

poisoning of downstream catalysts, was excessive, as was the need to replace

downstream catalysts poisoned by sulfur, and carbon was soon replaced by beds

of more reliable zinc oxide. Zinc oxide beds were well established after having

been used in hundreds of town gas, ammonia and hydrogen plants based on the

naphtha steam reforming process. Since it replaced activated carbon zinc oxide

has been almost exclusively used to protect steam reforming units.

9.3.2 Hydrodesulfurization

Zinc oxide can absorb simple sulfur compounds such as hydrogen sulfide and

mercaptans to form zinc sulfide with mercaptans however some carbon may also

be deposited. Other organic-sulfur compounds are not absorbed completely, and

thermal cracking of these compounds results in the deposition of carbon onto the

zinc oxide, thus reducing its absorption capacity. For these reasons, particularly

when refractory organic sulfur compounds such as thiophene or thioethers are

present in the feed, it is usual to add hydrogen and to include a bed of cobalt

molybdate catalyst to hydrogenolyse the sulfur compounds to hydrogen sulfide,

which can then be absorbed in the bed of zinc oxide. The cobalt/molybdate

component is sulfided during commissioning to the active form and then oper-

ates in the same way as in refinery hydrotreating. Although cobalt/molybdate

catalyst is the usual choice, nickel molybdate can also be used. This is important

when the hydrogen stream contains a high proportion of carbon oxides. Unsul-

fided cobalt/molybdate can catalyse the methanation of carbon oxides at temper-

atures higher than about 300°C, and this may lead to excessive bed temperature

while commissioning a new catalyst charge. This is less likely with the nickel

Synthesis Gas

359

catalyst. With liquid feeds, the catalyst can be presulfided at an appropriate low

temperature.

It is not usually necessary, however, to presulfide cobalt molybdate cata-

lysts before treating natural gas. Natural gas contains relatively small amounts of

simple sulfur compounds that are hydrogenolyzed at a low temperature. Cata-

lysts are sulfided during operation during operation and the eventual sulfur con-

tent of the catalyst depends on the sulfur content of the natural gas. Catalyst will

normally operate for several years with no loss of activity. The sulfur content of

typical feeds is shown in Table 9.4.

At the end of life in a typical ammonia plant, the cobalt/molybdate catalyst

will contain 0.5–3.0wt% sulfur and 2-10wt% carbon. The carbon content may

sometimes be higher than this but has no significant effect on performance. Hy-

drodesulfurization catalysts are usually pyrophoric after use because of adsorbed

hydrogen and carbon deposits. Before catalysts are discharged the reactor must,

therefore, be flushed with an inert gas such as nitrogen until the catalyst temper-

ature has fallen to less than 30°C. Large quantities of dust or carbon deposits on

the top of the catalyst layer can be sucked off and any broken or contaminated

pellets replaced with fresh catalyst. If the catalyst bed is to be completely re-

moved then a small volume of air can be added to the circulating nitrogen as the

catalyst cools. Adsorbed hydrogen or hydrocarbons are oxidized. This will lead

to a temperature increase and this can be controlled by limiting the volume of air

added. The catalyst can then be discharged at ambient temperature after cooling.

TABLE 9.4. Sulfur Content of Typical Steam Reformer Feedstocks.

Natural Gas North Sea Holland Associated Gas

Hydrocarbon/sulfur

(vol%) 81–94 81–82 70–75

(vol%) 4–5 3 12–15

(vol%) 0.1–0.5 0.5 5–8

+ (vol%) 0.08 0.2 5–9

(vol%) 0.2–0.5 1 Traces

(vol%) 0.8 14.3 –

Sulfur 5–30 ppm 1 ppm 1–3%

Naphtha

Specific gravity 0.71 0.74

Initial boiling point (°C)

33 43

Final boiling point (°C)

172 214

Total sulfur ppm (w/w) 230–260 1500

RSH 70–80 200–300

10–20 100–150

䩔

–

130

700–800

Average molecular weight ~100 ~110

360 Chapter 9

It is not economic to regenerate discharged catalysts for further use following a

typical life of several years.

9.3.3 Chlorine Removal

The presence of chlorides in the feed gas can cause the zinc oxide to sinter as

they are absorbed to form zinc chloride. Any hydrogen chloride formed in the

hydrodesulfurizer can be removed from the gas by installing a bed of alkalized

alumina in front of the zinc oxide. After this treatment the natural gas usually

contains less than 0.1mg chlorine per cubic meter. An average of about 10–12

wt% chloride can be absorbed before the alkalized alumina is saturated. Plat-

former hydrogen containing up to 2 ppm HCl is commonly treated in this way.

Only a relatively small bed of absorbent is required at a space velocity of

10,000–15,000 h

-1

9.3.4 Sulfur Absorption

Zinc oxide has been used to remove sulfur compounds from hydrocarbons since

the 1930s, when the steam reforming process was first introduced. When a zinc

oxide composition is specially prepared to have a high surface area and high

degree of porosity, it can absorb more than 20 wt% sulfur in a single bed. It was

the preferred choice in the early days, because the partial pressure of hydrogen

sulfide at equilibrium under reaction conditions, particularly in the presence of

water vapour, is very small compared to that of the bog iron ore used previously

to purify water gas/producer gas.

This observation was confirmed during the 1960s when luxmasse, a cheap

form of iron oxide, was investigated as an absorbent to desulfurize naphtha feed

in the British Gas Council Catalytic Rich Gas (CRG) process. Under typical

operating conditions of 370-400ºC and with a vapour phase water concentration

of 0.2–0.3%, the concentration of hydrogen sulfide in the vapour at equilibrium

is less than 3x10

-3

ppm when using zinc oxide, compared with 0.2–0.3 ppm

when using luxmasse. The data presented in Table 9.5 shows an increasing dif-

ferential as the concentration of water vapour increases. A further disadvantage

of luxmasse is that it is also reduced to magnetite before conversion to iron sul-

fide by reaction with hydrogen sulfide. Thus more water vapour is produced by

the conversion of luxmasse to iron sulfide, than is the case in the conversion of

zinc oxide to zinc sulfide:

3 Fe

+ H

Æ 2 Fe

+ H

O (9.4)

2 Fe

+ 8 H

S Æ 2 Fe

+ 8 H

O (9.5)

Synthesis Gas

361

TABLE 9.5. Equilibrium Hydrogen Sulfide Concentrations.

Water

(vol%)

Equilibrium Concentration H

S (ppm)

200°C 300°C 400°C

ZnO Fe

0.17 1.3x10

-5

1.57x10

-3

3.0x10

-4

9.3x10

-2

3.3x10

-3

1.96x10

-1

0.33 2.6x10

-5

8.5 x10

-3

7.0x10

-4

2.0x10

-1

6.5x10

-3

4.9x10

-1

1.7 1.3x10

-4

7.0 x10

-1

3.0x10

-3

2.0 3.2x10

-2

4.23

Overall:

3 Fe

+ H

+ 8 H

S Æ 2 Fe

+ 9 H

O (9.6)

Whereas:

6 ZnO + 6 H

S Æ 6 ZnS + 6 H

O (9.7)

Thus six atoms of iron will produce nine molecules of water compared with six

atoms of zinc, which only produce six. Furthermore iron sulfide can release hy-

drogen sulfide by reaction with hydrogen and steam during plant start-up or

shut-down conditions. In practice, therefore, when iron oxide is used it is always

followed by a guard bed of zinc oxide.

9.3.4.1 Operation with Zinc Oxide

In general, more sulfur is absorbed by zinc oxide at higher operating tempera-

tures, and when the porosity of the zinc oxide is increased. Consequently, oper-

ating temperatures are usually greater than 300ºC, and porous particles are pre-

ferred. However, the techniques needed to produce porous zinc oxide also lead

to a decrease in particle strength, which results in fracture of the particles during

loading and service, which, in turn leads to an increase in pressure drop across

the bed. Furthermore, there is also a decrease in the bulk density of the bed,

meaning that a lower mass of zinc oxide is charged to the reactor, and thus the

maximum amount of sulfur that can be absorbed by the bed also decreases pro-

portionally. This means that it is possible for a fixed bed size of highly porous

zinc oxide to absorb less sulfur than an equivalent bed of less porous, but more

densely packed zinc oxide. The porosity of granules must therefore be optimized

to give maximum absorbing efficiency and particle strength in any application

within the temperature range from ambient to 400°C. Typical operating condi-

tions are given in Table 9.6.

Up to about 10 ppm of low molecular weight mercaptans can be removed

from natural gas without hydrogen addition, provided that the reaction tempera-

362 Chapter 9

TABLE 9.6. Sulfur Absorbent Operation.

Operating Conditions Zinc Oxide Bed

Temperature (°C)

370–400

Pressure (atm) Up to 40

Space velocity (h

-1

) 500–1000

400°C 300°C 200°C

Sulfur capacity wt%S

(before slip observed)

25–30 20–25 8–10

By increasing the porosity of the absorbent the sulfur absorption can be increased by 200% in the

range ambient to 100°C and almost 100% at 200°C.

ture is high enough to crack the sulfur compound. Carbon deposition does not

immediately affect the sulfur absorption and a reasonable life can be achieved.

For a longer life, or with increasing concentrations of organic sulfur compounds,

a few percent of hydrogen is added to the gas stream being treated. This is rou-

tinely done in steam reforming plants which also nowadays include a bed of

cobalt molybdate catalyst. Hydrogen addition, in this case, is also beneficial

because it helps to reduce the reforming catalyst at the top of the steam reformer

tubes.

Carbonyl sulfide may be formed during natural gas treatment by the reac-

tion of hydrogen sulfide and carbon dioxide and this can be difficult to remove

completely using zinc oxide alone. This is not a problem if a bed of cobalt mo-

lybdate catalyst is included in the desulfurizer. It has been shown that lead oxide

is an efficient absorbent of carbonyl sulfide so the natural lead oxide impurity in

some zinc oxides may promote absorption, especially if traces of water vapour

are present to hydrolyze the carbonyl sulfide.

The operating temperature and space velocity through the beds of cobalt

molybdate and zinc oxide are chosen to give optimum performance for the sys-

tem as a whole. The maximum inlet temperature to the bed containing the co-

balt/molybdate is usually about 400ºC and the volume of zinc oxide is calculated

to give an acceptable period of time on line prior to discharge, for a hydrocarbon

feed with given sulfur content. This usually results in a space velocity of about

500-1000 hr

-1

in both beds and an operating life for the zinc oxide of at least one

year.

It can be more economical to have two beds of zinc oxide in series, with a

by-pass line, so that one bed can be changed with the other still operating. This

procedure allows an average of 25–30wt% sulfur pick-up during the period of

operation. With one bed the average pick-up will probably be about 20%. Zinc

oxide can desulfurize hydrocarbons at ambient temperature although the total

sulfur pick-up will only be about 5% before saturation.

Synthesis Gas

363

9.3.4.2 Preparation of Zinc Oxide

Zinc oxide absorbents are supplied in a variety of shapes such as spheres, extru-

dates, or pellets. These are generally made by granulation or extrusion of the

zinc oxide with small amounts of a suitable binding agent such as a cement. The

final zinc oxide content ranges from 90-100% and the bulk density from about

0.9–1.5 kg liter

-1

. High density catalysts do not necessarily have the largest ca-

pacity for sulfur because the pore volume can also be low and this limits absorp-

tion. An optimum balance between bulk density and pore volume is, therefore,

essential to ensure that the maximum amount of the zinc oxide in a particle is

fully used. This not only extends the useful operating life but also minimizes

operating costs. The importance of proprietary catalyst recipes and raw materi-

als to improve catalyst performance have been explained in Chapter 1. Typical

compositions are given in Table 9.7.

9.3.4.3 Desulfurization of Other Gases

Carbon dioxide may be desulfurized with zinc oxide at the same operating con-

ditions as for hydrocarbons. Below about 150°C, however, zinc carbonate is also

formed at temperatures below about 150°C but this is subsequently converted to

zinc sulfide until saturation is reached.

9.4 STEAM REFORMING

Hydrocarbons are converted into a mixture of hydrogen and oxides of carbon by

reaction with steam over steam reforming catalysts. The reforming reaction is

endothermic and the catalysts are packed into narrow tubes, which are heated in

a furnace. The reforming furnace is commonly known as a reformer. An effi-

cient methane steam reforming process was developed by 1936

6,7

and was first

used on a large scale in North America during World War Two as shown in Ta-

TABLE 9.7. Sulfur Absorbent Composition.

Property A B

Zinc oxide (wt%) 90 75–85

Calcium oxide (wt%) 3–4 –

Alumina (wt%) Balance 4–5

Silica (wt%) – 5–10

Ferric oxide (wt%) 1–2 –

Bulk density (kg liter

-1

) 1.0–1.1 1.0–1.1

Surface area (m

-1

) 25–35 20–25

Shape Granules Extrudates

364 Chapter 9

ble 9.8. Early reformers operated at slightly higher than atmospheric pressure.

As demand for fertilizer nitrogen increased, however, and better alloys for re-

former tubes were developed higher reforming pressures could be used in larger

ammonia plants.

The original raschig ring catalyst, which contained precipitated nickel ox-

ide, kaolin and a silica cement, was more or less unchanged until high-pressure

reformers were introduced. Some of the early catalyst charges to be used lasted

for about 20 years.

Eventually, improved catalysts were required in modern

plants because silica was found to be volatile in high-pressure steam and the

catalyst rings part way down the tubes became weakened. At first, silica free

catalysts were made by exactly the same procedure, simply excluding silica

from the ingredients and using alumina cements. However, all modern, high

activity, reforming catalysts are now based on preformed, thermally stable, sili-

ca-free supports impregnated with nickel oxide. Ring catalysts are still available

but the same formulation can be provided as special, high geometric surface area

shapes. Shapes require pressure drop through the tubes and are able to transfer

heat rapidly from the tube wall to the reacting gases. As a result of better heat

transfer and higher activity, lower gas temperatures are effective in the tubes and

the furnace.

Until 1960, steam reformers were only used in areas with readily available

supplies of natural gas. As the cost of handling large quantities of coal increased

during the 1950s and large volumes of cheap naphtha became available from

refineries, the range of feedstock that could be reformed was extended by the

introduction of a new alkalized reforming catalyst based on nickel. This began a

revolution in synthesis gas production throughout the world. This new catalyst

enabled the use of cheap naphtha with a boiling point up to 220C to be used,

without the deposition of carbon on the catalyst and furthermore, there were no

compression costs in the operation of the reformers at higher pressures. The de-

TABLE 9.8. The First US Steam Reforming Ammonia Plants.

a,b

Location Contractor

Ozark Ordnance Works, Eldorado, Arkansas. Chemico (four furnaces)

Jayhawk Ordnance Works, Baxter Springs, Kansas. Chemico (two furnaces)

Cactus Ordnance Works, Etter, Texas. Chemico (three furnaces)

Ammonia Plant, Sterlington, Louisiana. M W Kellogg (two furnaces)

Alberta Nitrogen Products Ltd, Calgary, Alberta. M W Kellogg (two furnaces)

All except Sterlington used ICI catalyst.

Each furnace contained 66 tubes

Chem Week Report (Ammonia), Sept. 11 (1965) 11. Other plants listed in article are based on coal

feed.

N Gard, Thirty Years of Steam Reforming – A Review of ICI Developments & Experience, Nitro-

gen, Jan/Feb 1966.

Figure

synthes

of M. T

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stream

9.4.1

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in a fu

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be su

9.1. Schematic

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Synt

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imary reforme

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the large sin

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nsured that t

ammonia on

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t, and suspen

9.1 and 9.2).

of reaction m

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366

Figure

former

ermis

theref

strengt

cause

the li

time.

Chapter 9

9.2. Schematic

in synthesis ga

ion of M. Twig

re have excep

during such

w-pressure re

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hysical

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arrangement o

production. R

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Synthesis Gas

367

would deform or creep at the high operating temperatures required for the steam

reforming process. The lifetime of the reactor tubes before failure was the main

limitation to progress, until more thermally resistant alloys became available.

When designing a reformer the effects of stress at the high operating tempera-

ture and pressure are taken into account to select the most economic tube mate-

rial for a life of about 100,000 hours.

For effective operation, heat must be transferred rapidly from the furnace it-

self to the surface of the catalyst, particularly at the top of the tubes. This is to

ensure that the hydrocarbon feed reacts rapidly with the steam, as this helps to

avoid the cracking reactions that lead to deactivation of the catalyst, and which

in turn, helps to avoid overheating of the reactor tubes. This is achieved by the

production of a catalyst with a high stable activity and by shaping the catalyst in

such a way that allows rapid mixing of the gases in the tube, to allow effective

heat transfer from the wall of the tube to the catalyst. This also results in a uni-

form temperature gradient throughout the tube. Careful packing of the catalyst

rings or shapes into the narrow, 3-4 inch diameter tubes is necessary to equalize

the gas flow and pressure drop through each tube. This also helps to avoid cata-

lyst breakage and voids within the catalyst bed. Any of these problems can lead

to maldistribution of gas flow, which causes variable tube temperatures and hot

spots.

Reformers built during the l940s, which operated at up to 50 psig, consisted

of about 66 tubes made of rolled and welded tubes made from type 310 stainless

steel. The early, high-pressure reformers built during 1965–1970 operated at up

to 500psig and contained up to 400 cast tubes made from HK40. Eventually,

when even better alloys like Manaurite 36X were introduced the HK40 tubes

were also replaced. Reforming capacity could be increased by up to 10–15% at

the same heat flux but with a lower wall temperature by putting a larger volume

of catalyst into Mauranite tubes of the same diameter as usual, but with thinner

walls. This was particularly useful when better catalysts, giving a lower pressure

drop and even better heat transfer, became available.

The hydrocarbon reforming reaction gives a mixture of carbon monoxide

and water which is close to equilibrium at the usual operating conditions. The

reaction between steam and the hydrocarbon gives a mixture of carbon monox-

ide and hydrogen that is close to thermodynamic equilibrium at the temperature

and pressure of the reactor.

+ H

O ↔ CO + 3 H

(9.8)

The operating steam to hydrocarbon ratio, or steam ratio, must, however, be

higher than the stoichiometric level to avoid carbon formation on the catalyst, by

cracking reactions, and to provide enough steam to operate the water gas shift

reaction later in the process.