E&O30O137

A New P r o c e s s for Co-production of Ammonia and

Methanol

Ahmed Soliman

Cairo University, Faculty of Engineering, chemical Engineering Department

Abstract:

A new process for co-production of ammonia and

methanol is proposed. The process involves the production

of synthesis gas by oxygen blown autothermal reformer

(ATR) at a pressure of 40-100 bars, a methanol synthesis

loop at a pressure of 50-100 bars and an ammonia synthesis

loop at a pressure of 200-300 bars. The oxygen required for

the ATR is supplied by an air separation plant.

The synthesis gases from the ATR are cooled and

compressed, in a first stage compression, to the required

methanol loop pressure. The purge stream from the methanol

loop is sent to an intermediate temperature shift converter

ITSC followed by a physical solvent CO z removal unit and

them purified in a pressure Swing Adsorber (PSA).

The purified hydrogen from the PSA together with the

almost pure nitrogen from the air separation plant are

recompressed, in a second stage compression, to the

TESCE, Vol. 30, No.2 <s> December 2004

required ammonia loop pressure. Ammonia condensation

takes place at high pressure by cooling water then by chilled

fluid cooled in the air separation plant.

1. Introduction:

Ammonia and Methanol are the most common liquid

chemical commodit ies . The world production of ammonia is

about 150 million tpa and the world production of methanol

is about 30 mill ion tpa.

Since they are commodit ies , the main factor that controls

the market share of the product is product price.

The product cost is the sum of the fixed cost and

operating cost. The operat ing cost is mainly the gas price

which is very competi t ive in Egypt and Arabic countries. As

a result the production cost of both ammonia and methanol

in the region is controlled by the fixed cost.

The fixed cost per ton products is minimized by

increasing the capacity and by maximizing product

integration in both side-stream direction (Ammonia /

Methanol Co- production process) and down-stream

direction (Nitrogen Value chain from ammonia and carbon

Value chain from methanol) .

TESCE, Vol. 30, No.2 <s> December 2(KM

Due to the fact that ammon ia and m e t h a n o l are twin

i n d u s t r i e s r e g a r d i n g the process c o n c e p t poin t of view. The

up — s t ream s y n t h e s i s gas uni t of the a m m o n i a p r o c e s s

(wh ich r e p r e s e n t s the major i ty of p lan t cos t ) is a lmos t

i d e n t i c a l to tha t of methanol . The p u r g e s t ream from the

m e t h a n o l loop is useful for the a m m o n i a p lan t . Also a very

i m p o r t a n t fact is that the capac i ty of bo th the a m m o n i a and

the m e t h a n o l p l a n t s are now l i m i t e d by the i r l oops and not

by the u p - s t r e a m syn thes i s gas uni t due to the d e v e l o p m e n t

a c h i e v e d in G T L p l a n t s .

The A m m o n i a me thano l c o - p r o d u c t i o n concep t i nvo lves

the c o n s t r u c t i o n of one in t eg ra t ed p l an t for both p r o d u c t s

c o n s i s t i n g of a la rge up-s t ream s y n t h e s i s gas uni t fo l lowed

by a d o w n - s t r e a m low-p res su re m e t h a n o l s y n t h e s i s loop then

a h i g h p r e s s u r e a m m o n i a syn the s i s l o o p .

2. Conventional Ammonia Technology:

A typical conventional ammonia plant consists of the following units:

• Synthesis gas production unit, which consists mainly of a

primary tubular steam methane reformer working at a pressure

of (25-40 bar), and air blown secondary reformer.

• Shift conversion unit, which consists of high temperature shift

converter (HTS) working at a temperature range of (300 - 500

°C) and a low temperature shift converter (LTS) working at

temperatures around 200 °C where most of the CO is converted

to C02.

TESCE, Vol. 30, No.2 <a> December 2004

• C02 removal unit, which is mainly an absorption stripping unit

working with a chemical solvent which consists mainly of

amino derivatives.

• Final purification, which consists mainly of a methanator to

convert the remaining CO into methane and a dryer to get rid of

undesired water vapor

• Ammonia synthesis loop consisting mainly of:

1. Two-stage compression unit to raise the synthesis

gas pressure to 100 - 200 bar

2. Multi - stage ammonia synthesis reactor

containing iron catalyst.

3. Ammonia refrigeration cycle.

The maximum capacity of a typical conventional ammonia plant is less

than 2500 tpd.

The typical energy consumption of a conventional ammonia plant is

about (27-30 GJ/t) NH3

The fixed cost required for a conventional ammonia plant with a

capacity of 1500 tpd is about 260-270 MUSD, which is equivalent to fixed

charges of more than 50 USD/ton ammonia.

Figure (1) shows a schematic presentation of conventional ammonia

technology

TESCE, Vol. 30, No.2 o> December 2004

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3. Convent ional Methanol Technology:

A typical conventional methanol plant consists mainly of the following unit:

• Synthesis gas production unit, which consists mainly of a primary

tubular steam methane reformer working at a pressure of (25-40

bar).

• Methanol synthesis loop consisting of:

1. Two-stage compression unit to raise the synthesis gas pressure

to ( 5 0 - 1 5 0 bar).

2. Adiabatic or quasi-isothermal methanol synthesis reactor

containing Copper oxide - Zinc oxide - Aluminum oxide

catalyst.

• Distillation train to purify crude methanol from light and heavy

ends.

The typical energy consumption of a conventional methanol plant is about

(28 - 32 GJ/t methanol).

The fixed cost required for a conventional methanol plant with a capacity

of 2500 tpd is about 300 MUSD, which is equivalent to fixed charges of more

than 30 USD/ton methanol.

Figure (2) shows a schematic presentation of conventional methanol

technology

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4. Modern Trends in S y n t h e s i s Gas Technology:

The continuous development in research in synthesis gas production

technologies, especially for GTL plants, and the development of new

catalysts, materials and new equipment lead to considerably gTeat impact on

the design of modern plants. The new developments include: autothermal

reformer & heat exchanger reformer.

4.1 Fully Autothermal Reforming:

The whole reforming reaction could be performed without a tubular

reformer by autothermal catalytic reforming in a design similar to a secondary

reformer. In- this case it would be necessary to use oxygen or oxygen-enriched

air instead of air.

The auto-thermal reformer is fed directly with hydrocarbon feedstock or

with partially reformed gas in a pre-reformer. Because of the higher heat of

reaction in the internal combustion (temperatures of 2000 °C and higher), the

flow conditions, heat release characteristics and the risk of soot formation are

very different from the situation in the normal secondary reformer, and special

considerations in the design of burner and reactor are necessary. The burner

must have excellent mixing characteristics which give a combustion zone with

a turbulent diffusion flame, followed by a thermal zone in which

homogeneous gas-phase reactions take place. In the following catalyst bed, a

close approach to the thermodynamic equilibrium in the methane reforming

TESCE,Vol.30,No.2 @ ) December 2004

reaction and the water gas shift reaction is attained. The combustion is a

substoichiometric reaction with an overall oxygen/ hydrocarbon ratio of

0.55—0.6.

Autothermal reforming of natural gas and also of naphtha at normal

pressure was used in large scale in BASF in the early 1960s for the production

of ammonia. The catalyst beds consisted of three layers: an inert material on

the top to prevent back flashing into the mixing chamber, a small intermediate

layer consisting of a noble metal catalyst as starter, and a bottom layer of

nickel catalyst for the reforming reaction. Two versions were in use, one of

which was performed with a flame reaction in the empty space above the

catalyst bed, while the other was used for naphtha. Feed and pre-mixed

oxygen directly entered the catalyst bed, where the reaction was initiated in

the noble catalyst layer.

4.2 Heat-Exchange Reforming:

The temperature of the flue gas from a traditional reformer is usually

higher than 1000 °C, and the process gas at the outlet of the secondary

reformer is also at around 1000 °C. From a thermodynamic point of view it is

waste of energy to use this high level heat simply for raising steam and

preheating process air for the secondary reformer. The boiling temperature in a

125 bar main boiler on the secondary reformer outlet is only 325 °C, and the

process air is usually preheated in the reformer flue gas duct. A concept which

completely avoids a fired primary reformer is the exchanger reformer, which

TESCL, \ ol. 3d, X„.2 <s> December 2004

with some simplification may be viewed as tubular heat exchanger with the

catalyst inside the tubes, which are heated by the hot secondary reformer

effluent flowing on the shell side.

In some designs the rubes may be open at the lower end, in which case the

gas flow on the shell side consists of a mixture of the exit gases from the

secondary reformer and from the reformer tubes. Commercially operating

designs are the GHR of ICI and the KRES of M. W. Kellogg.

In the ICI gas-heated reformer (GHR) the reformer tubes consist of an

outer scabbard tube with an open ended bayonet tube inside, and the annular

space between the tubes is filled with the reforming catalyst. The

steam/natural gas mixture enters the tubes via a double tube sheet and flows

downwards through the catalyst, and the reformed gas leaves through the

bayonet tubes. To enhance the heat transfer from the hotter secondary

reformer outlet which flows on the shell side, the scabbard tubes are finned

and surrounded by "sheath" tubes.

Quite recently ICI has come out with a modified design. The bayonet tubes

are replaced by normal tubes attached to a bottom tube sheet by a special seal

that allows some expansion. In this way the delicate double, tube sheet of the

GI-IR is avoided. The seal which prevents leakage of methane-rich gas to the

secondary reformer effluent flowing on the shell side has a^uniquc design

which is"subject to patent applications of ICI. The AGHR will allow a single-

TESCE, Vol. 30, No.2 v ' IK timber 2004

line concept for world-scale plants whereas with the GHR several parallel

units would be necessary in the case of large plants.

Because of the smaller size compared to a conventional fired reformer,

considerable investment savings can be achieved. To close the heat balance

between the exchanger reformer and secondary (autothermal) reformer, the

latter has to take on a higher reforming duty, which may be achieved by using

an over-stoichiometric amount of air or oxygen-enriched air. In some

configurations, such as the KRES concept of Kellogg, the exchanger reformer

is partially bypassed, part of the feed (70-75%) being fed directly into the

autothermal reformer. The overall S/C ratio is 3.0 to 4.0 and the oxygen

content in the enriched air is between 28 and 32 mo! %. The mechanical

design is relatively uncomplicated, and the pressure difference across the tube

walls is only 3.4 bar (35-40 bar in a fired primary reformer). In contrast to the

1C1 GHR there is only one tube sheet and the tubes are open at the tower end,

where the reformed gas mixes with the hot effluent of the secondary reformer.

The mixed gases flow upwards on the shell side, where baffles create a cross-

flow to improve heat transfer. For mixing of the two gas streams, a multi-hole

distributor is installed.

Similar concepts are available from other licensors and contractors. Braun

& Root (now KBR), is offering the Tandem Reformer a process developed and

commercially tested in the former USSR in a hydrogen plant equivalent to a

400 t/d ammonia plant. Topsoe also has an exchanger reformer design.

(341 \ TKSCE, Vol. 30. No.2 ^ ' December 2004

A concept developed by Uhde goes a step further in this direction:

exchanger reforming and subsequent noncatalytic partial oxidation, which

provides the reaction heat, are accommodated in a single vessel. This

combined autothermal reformer (CAR) design, the reformer is operated at 17

bars in a demonstration unit producing 13000 m3/h of synthesis gas. The

steam/natural gas feed is passed through the reforming catalyst in the tubes,

which are heated by the hot gas returning from the partial oxidation section

below the tubes. The intensive mixing of the tube effluent with the oxygen or

oxygen-enriched air in the partial oxidation zone (temperature about 1300 °C)

is achieved by a vortex-type flow pattern. The residual methane content is

governed by the amount of oxidant. For ammonia production all the feed will

be passed through the reforming tubes; in other applications (methanol, oxo

gas) a portion of feed can be fed directly to the oxidation chamber to attain a

higher CO content.

5. Modern Trends in Ammonia Technology:

The continuous development in research in ammonia technologies and the

development of new catalysts, materials and new equipment lead to

considerably great impact on the design of modern plants. The new

developments include: intermediate temperature shift converter and ruthenium

catalyst.

TESCE, Vol. 30, No.2 <5> December 2004

5.1 Intermediate - Temperature Shift Convertor (ITSC):

A relatively new process concept is the intermediate temperature shift

which performs the reaction in a single step. The catalyst is based on a copper

-zinc- alumina formulation and optimized for operating in a wider temperature

range (200-350 °C) than the standard LTS catalyst (190-275 °C). The reaction

heat can be removed by use of a tube-cooled reactor raising steam or heating

water for gas saturation to supply process steam in the reforming section

(Linde LAC, ICI Catalco LCA). In a new plant using the spiral-wound Linde

reactor, a methane slip of only 0.7 mol% (dry basis) is achieved. Further

purification is performed by PSA. Generally the shift conversion reactors have

an axial gas flow pattern, but recently radial gas flow configurations have

been chosen in some instances. The lower gas velocities result in reduced

pressure drop, which saves compression energy and allows the use of smaller

catalyst tablets (3 x 3 mm or 2 x 2 mm), which because of the higher activity

can achieve a-Iower CO leakage. In a 1000 t/d plant without hydrogen

recovery the ammonia production would be increased by 10 t/d. Ammonia

Casale has patented radial flow designs for both HTS and LTS. In a recent

revamp the pressure drop was reduced from 0.55 to 0.35 bar in the HTS and

from 0.55 to 0.22 in the LTS. The CO leakage from the LTS could be lowered

to 0.11% from 0.25%.

5.2 Commercial Ruthenium Catalysts:

Since the early days of industrial ammonia synthesis only minor

improvements have been achieved for the magnetite system: optimization of

TESCE, Vol. 30, No.2 <343>

December 2004

manufacturing procedures, promoter concentrations, and particle size to give

somewhat higher activity and longer service life.

The most notable development for the magnetite system was the

introduction of co It as an additional component by 1C1 in 1984. The cobalt-

enhanced catalyst formula was first used in an ammonia plant in Canada using

ICI - Catalco's AMV process (later also in other AMV license plants) and is

also successfully applied in I d ' s LCA plants in Severnside.

In 1979 BP disclosed to M.W. Kellogg a new catalyst composed of ruthenium

on a graphite support. In October 1990, after a ten-year test program. Kellogg

started the commercialization of the Kellogg Advanced Ammonia Process

(KAAP) using this catalyst, which is claimed to be 10—20 times as active as

the traditional iron catalyst. According to the patent the new catalyst- is

prepared by subliming ruthenium-carbonyl [Ru3(CO)i2] onto a carbon-

containing support which is impregnated with rubidium nitrate. The catalyst

has a considerably higher surface area than the conventional catalyst and,

according to the patent example; it should contain 5 wt% Ru and 10 wt% Rb.

Besides having a substantially higher volumetric activity, the promoted

ruthenium catalyst works best at a lower than stoichiometric H/N ratio of the

feed gas. It is also less susceptible to self-inhibition by NH3 and has excellent

low-pressure activity A diameter of 1.5-2.5 mm and a length of about 6-7 mm

are probably the dimensions of the particles of the commercial catalyst

TESCE, Vol. 30, No.2 <s> December 200-4

The potential for ruthenium to displace iron in new plants (several projects

are in progress of which two 1850 mtpd plants in Trinidad now have been

successfully commissioned) will depend on whether the benefits of its use are

sufficient to compensate the higher costs. In common with the iron catalyst it

will also be poisoned by oxygen compounds. Even with some further potential

improvements it seems unlikely to reach an activity level which is sufficiently

high at low temperature to allow operation of the ammonia synthesis loop at

the pressure level of the synthesis gas generation.

6. Modern TrendS in Methanol Technology:

The continuous development in methanol technologies and the

development of new catalysts, materials and new equipment led to

considerably great impact on the design of modern methanol plants. The new

developments include new reactor designs and low pressure catalysts.

6.1 Modern Reactor Designs:

Ammonia- Casale S.A. has developed a reactor that employs a combination

of axial and radial flow (mixed flow). This type of reactor initially developed

for ammonia plants is offered by Davy McKee in ICI license.

The Variobar reactor developed by Linde consists of a shcll-and-mbc

reactor coiled in several tiers, whose cooling tubes are embedded in the

catalyst packing. The reactor temperature is adjusted by water cooling As in

other processes, the heat of reaction is utilized to produce steam, which can be

TESCE, Vol. 30, No.2 •(34T>

December 2004

used, for example, to drive a turbine for the compressor as an energy source

for subsequent methanol distillation.

The Mitsubishi Gas Chemical (MGC) process uses a reactor with double-

walled tubes that are filled in the annular space with catalyst. The synthesis

gas first flows through the inner tube to heat it up and then, in countercurrent,

through the catalyst between the two tubes. The outer tubes are cooled by

water; Mitsubishi considers the main advantage of this process to be the high

conversion rate (ca.14% methanol in the reactor outlet).

6.2 Low Pressure Catalyst.:

All currently used low -pressure catalysts contain copper oxide and zinc

oxide with one or more stabilizing additives. Alumina, chromium oxide, or

mixed oxides of zinc and aluminum have proved suitable for this purpose.

Table 1 shows a summary of typical modern catalysts for low pressure

methanol svnthesis.

TESLE, Vol. 30, No.2 V ' December 2004

Table 1 Summary of typical modern catalysts for low pressure methanol

synthesis.

i Manufacturer IFP

Sud Chemie

Shell

ICI

BASF

Du Pont

United Catalysts

Haldor Topsoe

| Component Cu Zn Al Cu Zn Al Cu Zn

Rare earth oxides Cu Zn Al Cu Zn Al Cu Zn Al Cu Zn Al Cu Zn Cr

] Content, atom % 2 5 - 8 0 1 0 - 5 0 4 - 2 5

6 5 - 7 5 1 8 - 2 3 8 - 1 2

71 24 5

61 30 9

6 5 - 7 5 2 0 - 3 0 5 - 1 0

50 19 31 62 21 17 37 15 48

TESCE, Vol. 3(1, No.2 o> December 2004

7. Ammonia — methanol co-production:

The new technology for co-production of ammonia and methanol involves

the following stages:

l .Air separation plant to produce almost pure nitrogen and liquid

oxygen (90% min).

2. Preheating of natural gas, steam and oxygen to a temperature (500 -

600 °C).

3. Pre-reforming of natural gas in a heat exchanger pre-reformer heated

by the outlet hot gases from the OBATR at a temperature (900 - 1100

°C).

4. Auto - thermal reforming of the pre-reformed gases in an OBATR at a

pressure 40-100 bar

5. The hot synthesis gases from the OBATR at (900 - 1100 °C) are used

for heating in the heat exchanger pre-reformer.

6. The hot synthesis gases leaving the heat exchanger pre-reformer are

further cooled in a waste heat boiler generating high pressure steam.

7. The cooled synthesis gas is compressed to a pressure 50 - 100 bar in

a first stage turbo-compression unit.

8. The compressed synthesis gas is feed to a low pressure methanol

synthesis loop.

9. The purge stream from the methanol synthesis loop is preheated to a

temperature (200 - 350°C) then fed to an intermediate temperature

shift convector (ITSC).

TESCE, Vol. 30, No.2 <s> December 20(14

10. The shift gas from the ITSL is then cooled and fed to a physica1

solvent C02 removal unit to separate C02 (this stage is importarr.

only if urea is to be produced).

11. The shift gases are then purified in a multi-bed pressure swint;

adsorption unit (PSA) to separate almost pure hydrogen (usually 99.9

%).

12. The hydrogen from the PSA is compressed with almost pure

nitrogen, from the air separation plant, in a second stage turbo-

compressor unit to a pressure 200 - 300 bar.

13. The compressed gases are fed to a high pressure ammonia synthesis

loop where ammonia is condensed by cooling water followed by a

chilled fluid cooled in the air-separation plant.

Figure (3) shows a schematic presentation for the proposed

Ammonia-Methanol co-production process.

(349s) TESCE, Vol. 30, No.2 ^ / Decern!.* = 20<>4

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8. References:

1. Fertilizer Manual, Kluwer Academic Publishers, Dordrecht,

UN1DO, 1FDC Hardbound, March 1998.

2. Max Appl, "Ammonia Principles & Industrial Practice", W1LEY-

VCH, 1999. 3. Ullmann's Encyclopedia of Industrial Chemistry Release 2004,

7th Edition.

TESCE, Vol. 30, No.2 <S>

December 2004