Higman Chris Gasification (Газификация угля)

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230

Gasification

van Liere, J., Bakker, W. T., and Bolt, N. “Supporting Research on Construction Materials and

Gasification Slag.” Paper presented at VGB (Technische Vereinigung der Großkraftwerkstbe-

treieber) Conference, “Buggenum IGCC Demonstration Plant,” Maastricht, November 1993.

van Loon, W. “De vergassing van koolstof met zuurstof en stoom.” (The gasification of

carbon with oxygen and steam). Ph.D. diss., Delft University, 1952.

Visconty, G. J. “Method of Transporting Powder into Advanced Pressure Zone.” U.S. Patent

2,761,575, 1956.

Weigner, P., Martens, F., Uhlenberg, J., and Wolff, J. “Increased Flexibility of Shell Gasification

Plant.” Paper presented at IChemE Confernce, “Gasification: The Clean Choice for Carbon

Management,” Noordwijk, April 2002.

Weiss, M. M. “Selection of the Acid Gas Removal Process for IGCC Applications.” Paper pre-

sented at IChemE Conference, “Gasification Technology in Practice,” Milan, February 1997.

Wildermuth, E., Stark, H., Ebenhöch, F. L., Kühborth, B., Silver, J., and Rituper, R. “Iron

Compounds” In Ullmann’s Encyclopedia of Industrial Chemistry, 5th ed., vol. A14, p. 596.

Weinheim: VCH Verlagsgesellschaft, 1990.

Wilhelm, S. M. “Effect of Mercury on Engineering Materials Used in Ammonia Plants.”

Paper presented at AIChE Ammonia Plant Safety Symposium, San Diego, 1990.

231

Chapter 7

Applications

7.1 CHEMICALS

The two chief components of synthesis gas, hydrogen and carbon monoxide, are the

building blocks of what is often known as C1 chemistry. The range of products

immediately obtainable from synthesis gas extends from bulk chemicals like ammonia

and methanol, through industrial gases, to utilities such as clean fuel gas and

electricity. Furthermore, there are a number of interesting by-products, such as CO

and steam. As can be seen from Figure 7-1, many of these direct products are only

intermediates toward other products closer to the consumer market, such as acetates

and polyurethanes.

Synthesis gas is an intermediate that can be produced by gasification from a wide

range of feedstocks and can be turned into an equally wide range of products. And

although every combination of gasifier feed and end product is technically possible,

this does not mean that every combination makes economic or even technical sense.

In North Dakota, synthesis gas generated from coal is successfully processed to

manufacture synthetic natural gas (SNG). In Malaysia, partial oxidation of natural

gas is used to generate the synthesis gas feed for a synthetic liquid fuels operation.

Yet it would clearly make no sense to generate synthesis gas by partial oxidation of

natural gas to manufacture SNG.

Given that this broad range of products is available from the single intermediate

of synthesis gas, there is no technical reason why one could not produce more than

one product from the same gas source. In fact, many operators of gasification plants

do precisely this. This is known, in an analogy with co-generation (electricity and heat),

as polygeneration. Some even go a step further and install surplus downstream cap-

acity compared with the available syngas generation capacity. In this manner, such

operators are able to “swing” production from one product, say ammonia, to

another, say methanol, or peaking power in accordance with market demand and are

thus in a position to optimize revenue from the gasification plant. In a reverse manner,

there are other operators using different feedstocks, and even where appropriate differ-

ent technologies, to generate their syngas. In such a case, the opportunity is to work

with the cheapest feedstocks, topping up with more expensive ones only as required.

This inherent flexibility associated with syngas production and use provides a

multitude of choices that is increased by the variety of utility systems, in particular

232

Gasification

the broad possibilities for steam system configuration. It is therefore useful to look

at some typical gas processing designs for a number of the commoner applications

and review the considerations behind them.

7.1.1 Ammonia

Market

Over 90% of the world’s ammonia production capacity of 160 million t/y in 2001 is

based on steam reforming of natural gas or (in India) naphtha. Almost all the rest,

some 10 million t/y, is based on gasification of either coal or heavy oil.

The worldwide production of ammonia is by most measures the largest of any bulk

chemical. The principle use of ammonia is as nitrogenous fertilizer for agriculture.

Typical plant sizes today are 1500–2000 t/d. Process licensors are currently

revealing plans for plants up to 4000 or 5000 t/d size (Davey, Wurzel, and Filippi

2003; Parkinson 2001).

Synthesis Gas Specification

Ammonia synthesis takes place at high pressure over a catalyst that is usually iron,

although one process uses ruthenium according to the reaction:

AMMONIA

METHANOL

CARBON

MONOXIDE

GAS

TURBINES

REDUCTION

GAS

TOWN GAS

SNG

HYDROGEN

OXO

ALCOHOLS

FISCHER

TROPSCH

UREA

FORMAL-

DEHYDE

RESINS

MTBE

ACETIC

ACID

PHOSGENE

POLY-

URETHANE

DETERGENTS, PLASTICIZERS

FUELS, WAXES, OTHERS

COAL

LIQUID

RESIDUES

NATURAL GAS

BIOMASS

WASTE

METALS

ELECTRIC POWER

GASIFICATION

Figure 7-1.

Figure 7-1.Figure 7-1.

Figure 7-1. Applications for Synthesis Gas

Applications

233

+3 H

2NH

−92 MJ/kmol– N

(7-1)

A typical specification for ammonia synthesis gas is (Mundo and Weber 1982):

Ammonia plants based on gasification technologies normally surpass these specifi-

cations when a liquid nitrogen wash is the last stage of purification.

Most ammonia plants are built in conjunction with urea plants, the CO

from the

ammonia plant being used directly for urea production. According to the reaction

2 NH

+CO

CONH

O (7-2)

one mole of CO

is required for every two moles of ammonia. Typical requirements

for the CO

are as follows:

>98.5 mol%

S + COS <2 mg/Nm

<0.15mol%

Methanol <10 ppmv

Design Considerations

To process a raw synthesis gas to conform to this specification a number of different

tasks must be completed:

• Tar and volatiles removal (if present in raw gas)

• Desulfurization

• CO shift (CO +H

O = >H

+CO

)

• CO

removal

• Final removal of carbon oxides and water

• Adjustment of N

ratio

1:3 (For some modern processes nitrogen excess is

required)

CO + CO

<30 ppmv (As sum of total oxygen containing species)

O (in principle as for CO and CO

, but it can be

washed out with product ammonia in the synthesis

unit itself)

Sulfur <1 ppmv

Phosphorus, Arsenic,

Chlorine

(These are also poisons. Appl [1999] gives an

upper limit of 0.1 ppm for chlorine)

Inerts (including

methane)

<2%

minimum

←

→

←

→

234

Gasification

The order in which these tasks are performed may be modified to the extent that

there is a choice between performing the desulfurization before or after the CO shift.

Whereas older plants may have used a copper liquor wash for CO removal, most

modern gasification-based ammonia plants combine the tasks of CO removal and

nitrogen addition in a liquid nitrogen wash. In addition, the gas needs to be com-

pressed to the operating pressure of the synthesis, which can vary between 90 and

180 bar.

Prior to developing a block flow diagram, two (or perhaps three) key decisions

have to be taken.

First, a decision on the overall pressure profile of the plant needs to be determined.

In general, there is an energy advantage to running the gasifier at the highest possible

pressure, since the energy required to compress the synthesis gas is considerably

more than that required for compression of the feed materials. However, there are

three variables that one needs to consider:

•

The maximum pressure of the gasifier selected

. With oil feed, Texaco has numerous

references with 80 bar gasification. Shell can also operate SGP at this pressure,

but generally recommends operation at closer to 60 bar to reduce organic acid

formation in the system. For coal gasifiers, current commercial experience is limited

to 40 bar with one single unit at about 60 bar.

•

The oxygen compression system

. Here the choice is between compression of

gaseous oxygen and liquid oxygen pumping. Many 900 t/d or larger ammonia

plants operating with a gasifier at 60 bar have centrifugal compressors. Plants

significantly smaller than this would require the use of a reciprocating machine

with installed spare. Most of the 80 bar plants use liquid oxygen pumping. The

overall compression energy demand for liquid pumping (i.e., air plus oxygen

plus nitrogen) is some 3–7% higher than for the equivalent compression system.

This cancels some of the syngas compression advantages of an 80-bar system.

There are some safety arguments claimed in favor of liquid pumping, but the

excellent record of oxygen turbo-compressors in this service does not argue

against their use. All in all, there is not that much to choose between the two

systems.

•

The synthesis pressure

. In the 1970s and 1980s, typical operating pressures in the

ammonia synthesis were 220 bar. Today, most plants using conventional magnetite

catalyst are designed to operate at 130–150 bar, occasionally up to 180 bar.

Kellogg now uses a ruthenium catalyst operating at about 90 bar. The energy

demand for the synthesis does not change much over the range 130–180 bar, since

syngas compression gains made by operating at the lower pressure tend to be

compensated for by an increased refrigeration demand. Furthermore, at higher

pressures the volume of catalyst can be reduced, making the converter smaller

and, despite the increased design pressure, also cheaper. Often, therefore, the opti-

mization consists of selecting pressure levels for the gasifier and synthesis loop,

which minimize the number of casings on the synthesis gas compressor but

exploit that minimum number to the maximum extent.

Applications

235

• An additional factor to be aware of, even if it will not substantially influence the

choice of operating pressure, is that with increased pressures physical solvents for des-

ulfurization and CO

removal, gas losses through coabsorption of CO and H

will tend

to increase. In the range up to 80 bar, however, this remains within acceptable limits.

The second important decision is the selection of the acid-gas removal system and

its integration with the CO shift.

• When reviewing alternative gas treatment systems, the one immutable parameter

is the specification of the syngas. For ammonia production from gasification it is

not only a matter of eliminating carbon oxides, as is the case downstream of a

steam reformer. One needs to remove other components, some of which are

general to all gasification systems such as ammonia and HCN, others of which

may be feedstock or gasifier specific, such as the hydrocarbons produced by low-

temperature gasifiers. The system, which has proved itself capable of producing

on-spec gas behind practically all gasification processes, is Rectisol, which uses

cold methanol as a wash liquor. As a low-temperature process Rectisol is expensive.

However, in the ammonia environment this is not as serious as in other applications,

since a number of synergies can be achieved. All ammonia syntheses use some

refrigeration to condense the ammonia from the loop gas, and the product ammonia

is often stored in low-pressure tanks at a temperature of −33°C. The integration of

the refrigeration systems for Rectisol and ammonia synthesis allows some savings

compared with the stand-alone case.

Similarly, it is possible to integrate the refrigeration demand of Rectisol and the

liquid nitrogen wash, which operates at a temperature of −196°C. An additional

advantage of a physical wash is that CO

required for urea production can in part be

recovered under pressure, thus saving energy in CO

compression.

• One solution that is typical for use in conjunction with a syngas cooler is an imme-

diate desulfurization of the raw gas. After desulfurization the gas is “clean” but

also dry. In order to perform the CO shift reaction, it is necessary to saturate the

gas with water vapor in a saturator tower using water preheated by the exit gas of

the shift converter. CO

removal takes place in a separate step. The gas has thus to

be cooled and reheated twice in the process of acid-gas removal and CO shift. The

necessary heat-exchange equipment causes considerable expense and pressure

drop, a fact that has to be counted as a disadvantage of this system.

The raw gas from gasification of a typical refinery residue can contain about

1% H

S and say 4% CO

. A simple desulfurization of this raw gas will provide a

sour gas of about 20% H

S, sufficient for direct treatment in an oxygen-blown SRU.

Concentration of the H

S content up to 50% does not require much additional

expense. Furthermore, the CO

formed in the shift is free from sulfur and can be

used for urea production or emitted to the atmosphere without further treatment.

• When operating with a quench reactor, the gas emerges saturated with water vapor

at about 240°C. This temperature is too low to be able to generate high-pressure

steam, so it makes sense to utilize the water vapor immediately in a raw gas

236

Gasification

shift. In addition to saving the saturator tower, this has a number of minor advan-

tages, for example, that COS in the raw gas is also converted to H

S, and that HCN

is hydrogenated to ammonia (BASF undated). The sour gase components, H

S and

, are then removed all in a single step. Attractive as this appears, it is not with-

out its difficulties. In this case the “natural” sour gas has an H

S:CO

ratio of about

1:50. This requires considerable expense to concentrate the H

S to an acceptable

level for a Claus furnace and to clean the CO

before emission to the atmosphere.

Ultimately, however, there is not much to choose between these two systems.

A detailed study performed in 1971 (Becker etal.) compared a 60 bar scheme using

syngas cooling and clean gas shift with a 90 bar scheme using quench and raw gas shift.

It showed very little difference between the two routes. There was a slight energy

advantage for the syngas-cooler route, and the difference in investment amounted to

only 2%—lower than the estimating tolerance. This result continues to be valid to this

day, as can be judged by the commercial and operating success of both schemes.

Example

In order to elucidate these matters in more detail, we provide a worked example of a

1000 t/d ammonia plant based on heavy oil feed and using a syngas cooler and clean

gas shift. Based on the selected gasifier pressure of 70 bar and an overall pressure

drop of 15 bar over the gas treatment train, the syngas compressor suction would be

55 bar. This would allow a synthesis loop at 135 bar. For the sake of the example, it

is assumed that a liquid oxygen pump will be applied.

Before entering into a detailed description of the block diagram in Figure 7-2, there

are a number of further design considerations that need reviewing:

ASU CLAUS

POX

WASTE

WATER

TREATMENT

CARBON

HANDLING

REFRIGERATION

UNIT

AMMONIA

SYNTHESIS

LIQUID

NITROGEN

WASH

RECTISOL

CO SHIFT

RECTISOL

NITROGEN

SULFUR

AMMONIA

VANADIUM

CONCENTRATE

RESIDUE

WASTE

WATER

OXYGEN

CARBON SLURRY

AIR

Figure 7-2.

Figure 7-2.Figure 7-2.

Figure 7-2. Residue-Based Ammonia Plant

Applications

237

•

Oxygen quality

. Considering the fact that the final ammonia synthesis gas contains

25 mol% nitrogen, it is worth reviewing the extent to which this could and should

be brought into the system with the oxygen. The energy required for oxygen pro-

duction shows a flat minimum between about 90% and 95% purity. On the other

hand, an increase of nitrogen in the oxygen decreases the cold gas efficiency or

yield of H

+CO per kg residue. A detailed study will show an optimum at around

95% O

. This conclusion is also valid for an IGCC application where nitrogen in

the gas turbine burner is in any case required for NO

suppression. For many other

chemical applications (e.g., methanol), however, this would not be true, since N

is an unwanted inert in the synthesis loop.

•

Steam system

. The high-pressure steam must for safety reasons be capable of enter-

ing the gasification reactor under all conditions, and it is logical to generate steam

in the syngas cooler at the same selected pressure. For our example plant, we will

use 100 bar as the saturation pressure of the high-pressure steam and allow a

pressure drop of 8 bar across the superheater and controls. Provision will be made

for 25-bar medium-pressure steam and 10-bar low-pressure steam.

•

Compressor drivers

. With a syngas cooler installed after the gasifier, there is usually

sufficient high pressure steam to satisfy the demand of the CO shift, the syngas

compressor, and the nitrogen compressor. An external energy source is required

for the air compressor and (in this case) the nitrogen circulator required for oxygen

evaporation in the ASU. In the event of using an oxygen compressor, this would

substitute for the latter. An external energy source is also required for a refrigera-

tion compressor. Generally, two alternatives are available, electric power (pro-

vided the grid is stable enough to cope with starting a 12 MW electric motor) and

steam, which is generated in an auxiliary boiler on site. This is an economic

decision, but it should be borne in mind that start-up steam is required in any case.

Furthermore, having a substantial boiler in operation all the time can help in stabi-

lizing the overall steam system.

•

Refrigeration system

. The ammonia synthesis will require refrigeration capacity at

typically about 0°C. For final product cooling to atmospheric storage and for the

Rectisol acid-gas removal unit, additional refrigeration capacity at about −33°C is

required. Ammonia compression and absorption systems are available for this

duty. Absorption systems are generally more expensive in capital cost, but they

can operate on low-level waste heat that might otherwise remain unrecovered.

Studies performed in the context of ammonia plants like our example have shown

that it is more economical to use waste heat in boiler feed-water preheat than in

absorption refrigeration. On the other hand, if the waste heat really is “waste,”

then it pays to invest in an absorption system. Not only have both systems been

used in various locations, but they have also been built in combination, where an

absorption system was used to bring the refrigerant to 0°C and a booster compressor

was used for the −33°C duty.

Looking at the results of this discussion in Figure 7-2 together with the material

balance in Table 7-1 we see the following:

238

Gasification

Table 7-1

Mass Balance for 1000 t/d Ammonia Plant

Oxygen Steam

Raw

Gas

Desulf.

Gas

Sour

Gas

Shift

Gas Raw H

Pure

LNW

Syngas

, mol% 3.67 4.16 66.71 33.82 0.00 99.96

CO, mol% 49.28 49.46 4.37 3.20 4.74 0.02

, mol% 45.65 45.92 0.88 62.66 94.81 0.02 75.00

, mol% 0.27 0.27 0.08 0.19 0.26

, mol% 0.10 0.09 0.09 0.02 0.06 0.09 100.00 25.00

A, mol% 0.40 0.10 0.10 0.00 0.07 0.10

S, mol% 0.90 26.75

COS, mol% 0.04 1.19

, mol% 99.5

O, mol% 100.00

Dry gas (kmol/h) 1014 801 4250 4223 143 6116 3940 496 1321 4942

Pressure (bar) 65 72 55.1 52.5 1.5 49.5 47.6 1.5 50.0 44.5

Temp. (°C) 120 380 45 35 10 45 −60 20 45 35

Applications

239

Oxygen and nitrogen are manufactured in the ASU, where the compressors are

all driven by condensing steam turbines. The oxygen is pumped in the liquid

phase to a pressure of 80 bar and evaporated with gaseous nitrogen, which

returns the cryogenic cold to the cold box. The vacuum residue is gasified in

the partial oxidation reactor with oxygen and steam at 70 bar and about

1300°C. The raw gas from the reactor contains soot and ash, which is removed

in a water wash.

The raw gas, freed of solid matter is cooled down to about −30°C in the Rectisol

unit, where it is washed with cold methanol to a residual total sulfur content of

less than 100 ppbv. The sulfur-free gas is then heated up and saturated with water

at about 220°C in a saturator tower in the CO shift. Additional steam is added that

reacts over the catalyst with carbon monoxide to form hydrogen and CO

. The

gas at the outlet of the CO shift has a CO slip of about 3.2% and a CO

content of

about 34%. This gas reenters the Rectisol unit and is washed again with cold

methanol, this time at about −60°C. The CO

content is reduced to about

10 ppmv. The resulting gas is a raw hydrogen with about 92% H

and about 5%

CO, the rest being nitrogen, argon, and methane. This gas is cooled down to about

−196°C and washed with liquid nitrogen.

Simultaneously, the amount of nitrogen required for the ammonia synthesis is

added. The gas is then compressed to the pressure required for the synthesis loop.

The mass balance in Table 7-1, based on gasifying 32 t/h visbreaker residue in the

partial oxidation unit to produce 1000 t/d ammonia, illustrates this gas treatment

scheme. The feed quality is as described in Table 4-10.

7.1.2 Methanol

Market

Approximately 3.3 million metric tons per year, or about 9% of the estimated world

methanol production, is based on the gasification of coal or heavy residues.

Methanol is an important intermediate and, as can be seen from Figure 7-3,

approximately two-thirds of the production goes into the manufacture of formalde-

hyde and MTBE (methyl tertiary-butyl ether). The demand for methanol has varied

substantially from year to year, creating some dramatic price swings when supply

has failed to keep up with demand. At the time of this writing, the future of MTBE

as a component in reformulated gasoline is still uncertain, and this has its effects on

the market.

During the 1990s the typical size of world-scale natural gas–based plants

increased from 2000 to about 3000 mt/d. Current designs go up to 5000 t/d, and three

plants of this size are in the design stage (Göhna 1997). Most plants based on gasifi-

cation technologies are somewhat smaller. The largest, in the Leuna refinery in

Germany, has a nameplate capacity of 2060 mt/d.