Submission a Report (Version4)

SYNTHANOL SDN. BHD.

SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN

Table of Contents

1.0 Problem Definition..............................................................................................3

1.1 Feedstock and Product Specifications............................................................................3

1.2 Processing Objectives....................................................................................................4

1.3 Feedstock and Plant Availability....................................................................................4

1.4 Plant Capacity.................................................................................................................5

1.5 Scope of Design..............................................................................................................5

1.6 Definition of Terminal Points........................................................................................7

1.7 Site Characteristics and Constraints...............................................................................8

1.8 Utilities and Storages.....................................................................................................9

2.0 Technology Evaluation.......................................................................................11

2.1 Desulphurisation...........................................................................................................11

2.1.1 Gas-liquid contacting technology.........................................................................11

2.1.2 Solid bed Absorption............................................................................................14

2.1.3 Biological Process................................................................................................23

2.1.4 Selection of Technology.......................................................................................23

2.2 Syngas Production Technology....................................................................................25

2.2.1 Adiabatic Pre-reformer (APR).............................................................................26

2.2.2 Steam Methane Reforming (SMR)......................................................................28

2.2.3 Autothermal Reforming (ATR)............................................................................35

2.2.4 Combined Reforming...........................................................................................40

2.2.5 Heat Exchange Reforming...................................................................................41

2.2.6 Partial Oxidation (POX).......................................................................................43

2.2.7 Economics, Safety and Environmental Considerations for Reforming Process. .45

2.2.8 Selection of Reforming Technology....................................................................46

2.3 Methanol Synthesis......................................................................................................48

2.3.1 Three Phase / Slurry Phase Reactor.....................................................................50

2.3.2 Fixed Bed Reactor................................................................................................51

2.3.3 Adiabatic Quench Reactor....................................................................................52

2.3.4 Adiabatic Reactors in Series with Inter-stage Cooling........................................53

2.3.5 Tube Cooled Reactor............................................................................................55

2.3.6 Isothermal Boiling Water Reactor (BWR)...........................................................56

2.3.7 Reactor Selection..................................................................................................57

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2.3.8 Technology Evaluation of Catalyst......................................................................60

2.4 Product Purification......................................................................................................63

2.4.1 Single and Two-column Distillation Column......................................................63

2.4.2 Optimization of Process Technology...................................................................65

2.4.3 Selection of Process Technology.........................................................................75

3.0 Process Synthesis and Process Flowsheet Development.................................78

3.1 Development of Flowsheet Structure...........................................................................78

3.2 Reaction........................................................................................................................81

3.3 Separation.....................................................................................................................83

3.3.1 ATR Effluent........................................................................................................83

3.3.2 High Pressure Separator.......................................................................................83

3.3.3 Letdown vessel.....................................................................................................83

3.4 Recycle.........................................................................................................................84

3.4.1 Desulphurization Unit..........................................................................................84

3.4.2 Methanol Synthesis and Methanol Purification...................................................84

3.5 Overall Conversion and Yield......................................................................................85

3.5.1 Overall conversion...............................................................................................85

3.5.2 Yield.....................................................................................................................85

3.6 Economic, Safety and Environmental Consideration...................................................86

3.6.1 Economic..............................................................................................................86

3.6.2 Safety Considerations...........................................................................................89

3.6.3 Environmental Consideration...............................................................................90

3.7 Process Optimization....................................................................................................93

3.7.1 Steam Reformer....................................................................................................93

3.7.2 Autothermal Reforming.......................................................................................93

3.7.3 Methanol Synthesis..............................................................................................94

3.7.4 Methanol Purification...........................................................................................95

3.8 Process Flow Diagram.................................................................................................96

3.9 Process Flow with Reference to Process Flow Diagram..............................................99

3.10 Energy Integration......................................................................................................103

3.10.1 Heat Exchanger Network (HEN) Design...........................................................103

3.10.2 Process Flow Diagram With Heat Integration...................................................110

References....................................................................................................................115

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1.0 Problem Definition

1.1 Feedstock and Product Specifications

The feedstock for the methanol plant is natural gas (NG), which is sourced from

the natural gas reserves in South China Sea. Details about the sitting of this process

plant will be further discussed in Section 1.7. Table 1.1 below shows the natural gas

feedstock composition. This natural gas feedstock can be seen to be mostly consisted of

methane, but also have some amounts of ethane. Notably, it also has a small percentage

of hydrogen sulphide as well. The product of the plant, on the other hand, is methanol

and has specifications as described in Table 1.2. The minimum methanol content of the

product needs to be a minimum of 99.85 %.

Table 1.1: Natural gas feedstock composition.

Natural gas feedstock composition

Component mol%

Methane 88.73

Ethane 8.97

Nitrogen 0.45

Carbon dioxide 1.83

Hydrogen sulphide 0.02

Table 1.2: Methanol product specifications.

Methanol Specifications

Product properties Refined grade

Methanol content, wt% 99.85% min

Water content, wt% 0.15% max

Acidity (i.e. acetic acid), wt% 0.003%

Specific gravity (20 oC), g/cm3 0.7920 – 0.7930

Appearance Clear, no sediment

Permanganate number 30 min

Water miscibility No turbidity

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1.2 Processing Objectives

The main aim of this project is to design a processing plant capable of

converting the natural gas feedstock to methanol. This methanol will then be used in a

variety of applications including biodiesel production and fuel blending. It is crucial to

design and optimise the plant to produce methanol of required end specifications. All

factors including economics, environment as well as safety are to be taken into

consideration when designing the process. Waste effluents, as an example, will be

treated to meet environmental discharge standards before release into the environment.

1.3 Feedstock and Plant Availability

Natural gas should be readily available as feedstock prior to methanol

production. Kuantan Port City is chosen as the plant site which is located in Kuantan,

Pahang in Peninsula Malaysia This appears to be a strategic location for methanol

production due to the large amount of natural gas reserve in South China Sea, thus

enhancing the availability of feedstock obtained for the methanol production

(OECD/IEA, 2009). Moreover, the site in Kuantan Port City is chosen because of

several advantages as discussed in section 1.7. Figure 1.1 below shows the natural gas

pipeline infrastructure in South East Asia.

Figure 1.1: Natural gas infrastructure denoting feedstock availability (OECD/IEA, 2009).

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Pure Methanol(to storage tank)

Desulphuriser Pre-reformer Steam Reformer

Autothermal Reformer

Methanol Converter

Flash Separator

Refining Column

Feedstock Natural gas

1.4 Plant Capacity

The methanol plant is to be designed for a production capacity of 106 metric tons

per year of methanol. The plant is to operate for 330 days per year for 25 years. In this

design, a natural gas feedstock of 1925.19 mtpd is used to produce 3450.71 mtpd of

methanol. An overall carbon balance about the entire plant was done in order to justify

all mass balance calculations performed. This is clearly shown in Figure 1.2 below. A

total of 58590 kg/hr of carbon enters and leaves the system boundary. The required

amount of methanol produced at normal operation is 3030 mtpd. However, the plant is

designed for a capacity of 3450.71 mtpd of methanol. This is because in case of any

unforeseen circumstances, the demand of methanol can still be met by increasing the

capacity of the plant.

1.5 Scope of Design

The scope of the design will encompass the production of methanol (99.85%

purity) from feedstock natural gas. The general flowsheet for methanol production is

shown below in Figure 1.2.

Figure 1.2: Flowsheet of methanol production.

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mC , Me=53864 kg /hr

mC , Pur=4475 kg/hr

mC , AceAcid=251 kg /hr

SYNTHANOL SDN. BHD.


The first part of the project includes technology selection and evaluation for

each section. Different new and conventional technologies will be compared and the

more suitable and advantageous one will be selected for the plant design. Several factors

such as cost, operating conditions and lifespan will be taken into consideration.

Following this is the process flow diagram drawing which will show all the equipments

in correct sequence linked by pipelines. Required heating, cooling, compression and

pumping as well as utilities stream will also be shown. After making the required

assumptions and selecting the correct system boundaries, the material and energy

balance will be carried out to find out important parameters such as conversion, yield,

amount of recycle, waste gases and byproduct as well as all the heat loads and utilities

flow rates.

Several precautions will be taken in order to ensure no harm is done to the

environment. Air emissions shall be controlled to ensure the concentrations do not

exceed the discharge limits. Carbon capture will be practiced to ensure minimal

emission of greenhouse gases to the atmosphere. Any waste water produced will be sent

to a waste water treatment facility. Solid waste generated will be safely disposed

ensuring that no harmful substances are released during disposal. Within the plant,

catalyst regeneration will be carried out where possible. Sustainable practices will

include heat recovery as well as water reuse. In terms of safety, a hazard and operability

study (HAZOP) will be carried out to identify possible risks of explosion, fire, leakage

and collapse.

A draft Piping and Instrumentation Diagram (P&ID) will be drawn and after the

HAZOP the P&ID will be finalized. Control instruments will be installed to ensure

proper flow rates to the equipments. Alarms will also be installed to alert and correct

any deviations from required operating conditions.

As for the mechanical design section, each major equipment will be sized.

Appropriate equations and correlations for wall thickness, volume and pressure drop

will be used to design each vessel and piping from selected materials of construction.

Mechanical drawings will be constructed to scale for top and side views showing

correct orientation and relative size. A proper plant layout will also be drawn to show

the different positions of each section in the plant.

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The next step will be the economic evaluation of the project. The cost of all

equipments, raw materials, utilities and labour will be included for the calculation of

capital cost, operating cost and working capital. After taking into consideration tax

allowances and sales, the net present value will be calculated. The payback period and

internal rate of return will also be estimated.

In the end, based on the economic analysis and environmental considerations,

the project viability will be assessed.

1.6 Definition of Terminal Points

The processing line starts with the desulphurization of natural gas where the

natural gas is first preheated and compressed before entering the desulphuriser reactor.

The end point of the plant is a methanol storage tank which comes after the refining

column.

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Figure 1.3: Road Distance between Plant Site and Kuantan Port (Google Map, 2011).

SYNTHANOL SDN. BHD.


1.7 Site Characteristics and Constraints

The plant site is decided to be built in the Gebeng Industrial Estate in Kuantan

Port City with the coordinate of 3.983751, 103.381540 which is far from the residential

areas. Disturbance to the local community can be eliminated directly and it complies

with part of the amenity license condition (EPA Victoria, 2011). From Figure 1.3, the

site location is near to the Kuantan Port with the distance of 12 km through road access

in which the transportation expense is optimized to be the least from site to sea port for

importation and exportation of materials. The distance between proposed plant location

to Kuala Lumpur and Port Klang are estimated to be 263 km and 300 km respectively

(Figure 1.4 and Figure 1.5). Kuantan Port also offers trading possibilities to many

countries around the world. This will enable to conquer greater global methanol markets

to achieve higher profits as well as import raw materials from countries offering better

prices. Last but not least, Kuantan Port is a free-trade zone in which there are no taxes

accompanying with the material importation and exportation within the area. An

investment tax allowance of 100% qualifying capital expenditure for 5 years provided

by Malaysian government benefits the site in terms of economic consideration (East

Coast Economic Region Malaysia, 2010).

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Figure 1.4: Road Distance between Kuantan Port and Kuala Lumpur (Google Map, 2011).

SYNTHANOL SDN. BHD.


However, one constraint imposed on the chosen site is the unavailability of

railway. Railway tracks are important in the sense that tankers can be carried to and

from the plant in a more practical way. But, the Malaysian government has invested into

the development of a high-speed railway and inter-modal freight system and the

implementation of the project is already under way (Kaur, 2009).

1.8 Utilities and Storages

The utilities available to the plant are described in Table 1.3 below. The

associated cost per unit is also mentioned.

Table 1.3: Utilities available and the associated cost.

Utilities Cost of supply / treatment

Electricity: 11kV/3.3kV/415V 3Ph 50 Hz RM0.28/kWh

Natural gas: LHV 34.6 MJ/m3 (30 bar) RM600/t

Cooling water RM1.7/m3

Oxygen: Dry at 30 bar RM25/t

Saturated steam (30 bar) RM100/t

Hot water @ 90oC RM17.5/t

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Figure 1.5: Road Distance between Kuantan Port and Port Klang (Google Map, 2011).

SYNTHANOL SDN. BHD.


In terms of storage, the plant will have around 4 – 6 weeks of stocks of raw

materials and products which will constitute the working capital. Loading and

unloading facilities is required after the product storage tank.

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2.0 Technology Evaluation

2.1 Desulphurisation

Numerous processes have been developed to remove the sulfur content in the

natural gas based on a variety of chemical, physical and biological principles (Hairmour

et al., 2005).

Table 2.4: Typical feed gas specifications (Petersen et al., 2004).

ComponentsNatural Gas Associated Gas

Lean Heavy Lean Heavy

N2, vol% 3.97 3.66 0.83 0.79

CO2, vol% - - 1.61 1.50

CH4, vol% 95.70 87.86 89.64 84.84

C2H6, vol% 0.33 5.26 7.27 6.64

C3+, vol% - 3.22 0.65 6.23

Max total S, vol ppm 20 20 4 4

Hydrogen sulphide, vol ppm (typical) 4 4 3 3

COS, vol ppm (typical) 2 2 n.a n.a

Mercaptans, vol ppm (typical) 14 14 1 1

Sulfur scavenging processes can generally be categorized into solid bed (dry)

absorption process and liquid phase absorption process. Absorption of H2S into a liquid

occurs physically whereas chemical means of H2S removal involves adsorption of H2S

on a solid and further conversion into other sulfur-containing products (Pipatmanomai

et al., 2009). Biological conversion of H2S into elemental sulfur is possible by using

sulfide oxidizing microorganisms along with air or oxygen addition (Pipatmanomai et

al., 2009). The afore-mentioned H2S removal methods can be categorized into direct

stripping or direct oxidation (Hairmour et al., 2005). Besides that, there are also

available technologies to convert H2S directly to sulfur which are known as gas-liquid

contacting technology, such as Claus and LOCAT process.

2.1.1 Gas-liquid contacting technology

The gas-liquid contacting technology involves the use of a solution to either

scrub H2S from gas stream or strip H2S from liquid mixture. Claus technology applies

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the same working principle in which H2S is being absorbed into a solution and

regeneration can be carried out using air and subsequently form elemental sulfur (Smith,

2007). This technology has been successfully commercialized in different scopes of

industry in USA, Australia and Canada (Smith, 2007). CO2 is not absorbed in this

technology (Smith, 2007) and this is an advantage to the methanol synthesis process

since maximum methanol yield will be obtained provided none of the carbon-containing

compound is being depleted during the desulphurization process. However, there are

issues associated with wastewater problem and significant high capital cost involved

when liquid-based removal process is used (Pipatmanomai et al., 2009). Fouling

problems are possible to occur in Claus process (Smith, 2007) causing decrease in

process efficiency which will subsequently increase the unit operating cost.

Furthermore, Claus system was first introduced with the aim of treating tail gases from

various industries before releasing odorant H2S into the atmosphere hence it is not

suitable for natural gas treating (Nagl, 2007). Besides that, this technology is not

suitable to be employed for gas streams treatment with lower than 15% H 2S in the feed

stream (Nagl, 2007).

Hydrogen sulfide removal using liquid redox is another possible gas-liquid

contacting technology to be practiced in natural gas purification process. The state-of-

art in liquid redox technology is LOCAT® provided by the Gas Technology Product,

which uses an aqueous-based solution containing metal ions to carry out the redox

reaction. A non-toxic, chelated iron catalyst is used in this technology to accelerate the

H2S removal and subsequently forming elemental sulfur, which has economical value

(Nagl, 2007). The reaction involved is shown as below (Nagl, 2007):

H 2 S+12

O2

⇌ S+ H 2O

The H2S removal efficiency was reported to reach 99.9% and the operation can

be carried out at ambient temperature (Nagl, 2007). Surfside Environmental Inc.

(Removing Hydrogen Sulfide from Natural Gas Wells, 2011), which provides similar

technology using iron-base solution as well, also reported similar process features as

claimed by Merichem (Nagl, 2007). Due to the regeneration ability of the scrubbing

solution (Removing Hydrogen Sulfide from Natural Gas Wells, 2011) which is the iron-

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based aqueous solution, this technology provides a low operating cost as low as $0.13-

$0.16/ pound of sulfur removal (Nagl, 2007). The by-products generated from this

process are biodegradable liquid and sulfur in either molten or solid form, which has

economical value (Nagl, 2007). Besides that, liquid redox technology is capable of

handling any fluctuations in upstream compositions (Removing Hydrogen Sulfide from

Natural Gas Wells, 2011) and sulfur removing capacity as high as 20 tonnes per day

(Nagl, 2007).

Figure 2.7: Schematic representation of H2S removal unit using an ionized aqueous

medium (Removing Hydrogen Sulfide from Natural Gas Wells, 2011).

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Figure 2.6: A conventional H2S removal unit using liquid redox technology (Heguy et al., 2003).

SYNTHANOL SDN. BHD.


2.1.2 Solid bed Absorption

A fixed bed of solid particle can be used to remove H2S rather through chemical

reactions of physical ionic bonding. Typically, the natural gas stream must flow through

a fixed bed of solid particles which are known as catalyst or sorbents that remove sulfur

components and hold them in the bed. When the bed is exhausted, the media must be

replaced or regenerated. There are few commonly used processes under this category,

namely the zinc oxide process, iron-based scavenger process and adsorbents.

(i) Zinc Oxide (ZnO)

Among the adsorbents available in the H2S removal technologies, ZnO is known

to be a commodity sorbent and its reaction kinetics are well-studied (Sayyadnejad et al.,

2008). Besides that, due to the fact that ZnO has been used as either catalyst or sorbent

for the past 30 years in natural gas purification industry, its absorption capacity with

respect to different operating conditions could be easily predicted (Sayyadnejad et al.,

2008). The market for ZnO has been long established and it is readily available as

compared to any other sorbents (Sayyadnejad et al., 2008). The reaction mechanisms of

ZnO in H2S removal are shown below (Alphtekin, 2006):

ZnO+H 2 S⇌ZnS+H 2 O ∆ G=−17.5 kcal T=300℃

ZnS+2 O2⇌ZnSO4 ∆G=−16.3 kcal T=400℃

ZnSO4⇌ZnO+SO2 ∆ G=−6.7 kcalT=1000℃

When the natural gas is fed and passing through the catalyst bed made up of

ZnO, the traces of H2S existed will be absorbed by the active ZnO particles within the

catalyst. This is followed by the commencement of reaction at the outer surface of ZnO

particle which will eventually proceed to the core (Engelhard Corporation, 2005).

ZnO has the ability of absorbing both CO2 and H2S (Petersen et al., 2004). This

is considered as a disadvantage of using ZnO sorbent since maximum concentration of

CO2 is preferred in order to achieve higher methanol yield in the subsequent unit

operations. In general, there are at least two packed bed absorbers equipped in the

desulphurization unit in order to carry out the swing operation (Petersen et al., 2004).

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As the operating temperature decreases, the absorption capacity of ZnO towards H2S

will decrease as well (Atimatay et al., n.d.; Hairmour et al., 2005; Petersen et al., 2004).

The maximum absorption capacity per volume of ZnO installed will be limited by the

achievable bulk density of ZnO (Petersen et al., 2004). Engelhard Corporation, 2005,

showed that for a single bed operation, the maximum H2S absorption could be up to 30

wt% before saturation provided the sorbent used contained at least 95 wt% of ZnO.

There are factors affecting the H2S breakthrough capacity using ZnO as the sorbent. For

one, the presence of CO and its concentration will have an inverse effect on the

breakthrough capacity (Li, 2010). The operating temperature, H2O partial pressure and

the structure of ZnS formed will influence the breakthrough capacity as well (Li, 2010).

A pure ZnO sorbent will have structural changes at relatively high temperature

(Atimatay, 2008). This sintering effect will cause a decrease in the surface area

available for reaction to carry out and the shrinking effect of ZnO particles will increase

in severity as the operating temperature being further increased (Atimatay, 2008). Apart

from that, ZnO will tend to form metallic zinc vapor at temperature higher than 750℃

and hence limiting the maximum operating temperature of the ZnO-H2S system to only

750℃ (Atimatay, 2008). The lifespan of majority ZnO-based sorbents is limited

especially when ZnO sorbents are applied in fixed and moving bed reactors (Robert,

1994). The product formed after adsorption on ZnO, which is ZnS, has a molar volume

50% larger than that of ZnO. In other words, when zinc sulfate, ZnSO4 is formed during

regeneration, more than 250% of the volume originally occupied by ZnO is now

occupied by ZnSO4 (Robert, 1994). ZnSO4 formed during regeneration will somehow

decrease the reactivity of ZnO by blocking the catalyst pores as a result of its large

particle size (Karim, 2010). The repetition of continuous expansion and contraction of

sorbent due to the adsorption and regeneration cycle will instigate sorbent spalling. This

is a situation in which the sorbents will start breaking into smaller pieces and eventually

loses its function (Robert, 1994). In order to overcome this problem, fresh sorbent is

continuously supplied to the process and hence increasing the operating cost (Robert,

1994). In this case, sorbent with better durability will be a better choice to lower the

operating cost.

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Many literatures reviewed ZnO as a non-regenerable sorbent or difficult to

regenerate (Alphtekin, 2006). However, some research papers do categorize ZnO as

regenerable sorbent at particular regeneration temperature (Robert, 1994). Karim (2010)

reported on the possibility of regenerating the spent ZnO catalyst using two different

methods: combustion by air and steam treating of catalyst. Regenerated ZnO catalyst

was reported to be able to achieve activity as high as 97% and its physical and chemical

properties were comparable to commercial virgin ZnO catalyst (Karim, 2010). In recent

years, ZnO nanoparticles have been commercially produced to cater for the needs in

chemical industry. These ZnO nanoparticles with size ranging from 14 – 25nm emerge

to be a more effective H2S scavengers as compared to bulk ZnO particles used

commercially (Sayyadnejad, 2008). Spent ZnO sorbent is safe to dispose to the

environment without causing any adverse effect (Sayyadnejad, 2008). However, the

operating cost in the desulphurization unit will escalate as well if ZnO sorbent in

nanoparticles form is used instead of bulk ZnO. For sorbent consists of 100% ZnO, the

theoretical adsorption capacity was reported as 41.7 kg H2S/100 kg catalyst (Karim,

2010).

(ii) Zeolite Molecular Sieve

Zeolite molecular sieves are crystallized solids with very small evenly sized

pores. There are a large number of localized polar charges within the pores of the

crystalline structure which is known as the active site. The polar component in natural

gas such as H2S and water will enter the pores and form a weak ionic bonds at the active

sites, thus the H2S and water component will be trapped in the sieve. Cu (I) Y Zeolite

(Zeolite-Y) developed manage to reduce the sulfur content from 430 ppm weight of

sulfur to less than 0.1 ppm weight of sulfur This sorbents showed 40 times higher sulfur

selectivity and adsorption capacity as compared to other commercialized or

conventional sorbents due to formation of stronger bonds with hydrogen sulfide (H2S)

and other sulfur odorant molecules. A comparison study of the sulfur adsorption

capacity of cuprous zeolite with the other sorbents is conducted in a fixed bed adsorber

to evaluate the interaction between the sulfur compounds with the sorbents. The results

showed that Cu (I) Y zeolite possessed the highest adsorption capacities as compared to

AgY and Cu (II) Y zeolites. Besides that, Cu (I) Y zeolite displayed superior

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performance by being fully regenerable and showed least affinity for hydrocarbon

compounds (Crespo, 2008).

However, there are few disadvantages of using zeolite as sorbent for

desulfurization of natural gas. The aromatics or hydrocarbons compounds presence in

the raw natural gas is likely to compete for adsorption sites and hence reduce the

available active sites for sulfur adsorption. Furthermore, competition of the adsorption

site by water vapor will significantly reduce the sulfur capacity of the sorbents. Other

than that, carbon dioxide molecules are about the same size as H2S molecules, even

though CO2 is non-polar however the CO2 will still enter the pores and obstruct the

access of H2S to active site. Interference and competition for the active site of other

component is the main drawback of this technology. Nonetheless, molecular sieve using

zeolite is generally limited to small gas streams operating at moderate pressures. Due to

operating limitations, this technology is not commonly used for H2S removing

operations.

(iii) Iron-based Scavenger

Solid scavenger consists of iron-based materials which is able to remove H2S from

any gas streams (Nagl, 2007). Back in years ago, “iron-sponge”, a hydrated ferric-oxide

impregnated on wood chips, is used to carry out the following reaction, which is able to

convert H2S into some pyrophoric product (Nagl, 2007):

2 Fe2O3+6 H 2 S →2 Fe2 S3+6 H 2O

This process is applied to gases with low HsS concentrations (300 ppm) operating at low

to moderate pressures (50 – 500 psig). The Fe2S3 can be further oxidized with air to

product sulfur and regenerate the ferric oxide. However, the regeneration step involves

highly exothermic reaction with oxygen which possesses possibility to cause the wood

media to catch fire (Nagl, 2007). Furthermore, the bed has to be replaced after 10 cycles

of regeneration which induces highly operating cost.

The main disadvantage of this technology is such that the iron sponge media

often coated by the hydrocarbon liquids in the gas and hence inhibit the reactions. In

addition, the bed will eventually coat with elemental sulfur due to difficulty of

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controlling the regeneration step. Iron sponge units are normally operated in batch

mode, in order to achieve continuous production of methanol, a few iron sponge treating

unit are required. In economical point of view, this technology is not viable as compared

to the others. Moreover, the spent bed will continuously react with the oxygen in air

unless it is kept moist. Special handling of the waste is required which will indirectly

increase the operating cost of the plant.

(iv) SulfatreatTM Sorbents

The SulfatreatTM process offered by The Sulfatreat Co. at Chesterfield

implemented the Direct Oxidation Technology licensed from TDA Research

Incorporate for removal of sulfur. The Direct oxidation process catalytically converts

hydrogen sulfide (H2S) to elemental sulfur (S) and water (H2O) at 149 – 260 ℃. The

SulfatreatTM process is similar to iron sponge process. However, the iron oxides are

supported on the surface of an inert while the ferric oxide for iron sponge process is

impregnated on wood chips. This process is capable to achieve a 90% conversion of

H2S into elemental sulfur in a single pass (Jategaonkar, 2005).

Fex O y(s)+H 2 S(g )→

FeS2(s)+H 2 O(g)

According to Kohl et al. (1997), SulfatreatTM is composed of proprietary iron

compound, known as ferric oxide (Fe2O3) and ferrosoferric oxide (Fe3O4) which is

mixed with supplemental chemicals to produce a mixture of iron sulfides when react

with H2S. The conversion efficiency in commercial operations has been found out to

range between 0.55 and 0.716 lb H2S reacted / lb of iron oxide (Samuels, 1990) which is

somewhat higher than that of iron sponge bed design. Based on Samuels (1990),

significant improvement in operation and economics is observed by replacing iron

sponge with SulfatreatTM sorbents.

There are few advantages of implementing SulfatreatTM sorbents which make it a

potential technology to replace the conventional ones. One of the advantages of

SulfatreatTM is uniform porosity causes low pressure drop across the sorbent bed without

gas channeling. On top of that, uniform porosity and permeability of the sorbent only

allows reaction with sulfur-containing compounds. In other words, it hinders side

reactions with carbon dioxide (CO2) and other compounds. Besides, it is a non-

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pyrophoric substance hence eliminating the risk of fire. The starting material and spent

product are both safe and stable, where the spent product can be recycled or disposed

directly to landfill without any need of special handling. Based on a relative Screening

Index (SI) study by Foral et al (1993) which took into account investment and operating

cost, subjective weightings of process reliability, ease of operation, operator acceptance,

ease of spent material disposal and winterization requirements, SulfatreatTM has the best

rating and lowest total plant investment as compared to other technologies such as iron

sponge, zinc oxide, activated carbon, nitrite-based, SulfaRid and so forth.

The most basic equipment design for H2S removal with SulfatreatTM is a single

packed bed vessel operates in batch manner. For single vessel process, the natural gas

has to pass through a separator to eliminate large particulate prior to being fed into the

packed bed. The concentration of H2S at the outlet is only able to achieve non-

detectable levels at the beginning of the bed life, where the removal efficiency decreases

rapidly over time. Replacement of SulfatreatTM media is necessary once the outlet

concentration of H2S exceeded the specification level. The main drawback of this

system is that replacement of SulfatreatTM media requires temporary bypass of the vessel

which will directly interrupt the process flow. Hence, this system is not favourable to

meet the need of continuous production of methanol.

Therefore, a Lead/lag arrangement is chosen in which two vessels are arranged

in series as shown in Figure 3. This is to increase the efficiency of the SulfatreatTM

sorbents with no interruption in unit service and enhances the process reliability.

Lead/lag vessels are able to improve the overall removal efficiency of the system as

high as 20%. All the H2S will be removed at the beginning of the treatment when the

flowing gas passed through the first vessel which acts as the “working” unit. The exit

gas will enter the second vessel, “lag” unit, for further purification when the level of

outlet H2S reaches the maximum specification or act as a backup working unit. The

SulfatreatTM material is considered spent or exhausted once the inlet and outlet

concentrations reaches unity with typical lifespan of two to three years. Then, the

second vessel will be the lead unit whilst spent sorbents will be replaced with fresh

SulfatreatTM without interrupting the flow (Mi Swaco, 2010).

19

SYNTHANOL SDN. BHD.


Figure 2.8: Lead/Lag system of Sulfatreat process (SulfaTreat, 2011).

(v) Sulfatrap™ Sorbents

Recently, TDA Research Incorporate has developed SulfatrapTM series of

sorbents to effectively remove sulfur from natural gas and has been carrying out field

demonstration for the past two years over the United States of America. SulfatrapTM-R2

and R6 series are developed particularly for natural gas desulfurization process with

high selectivity and sulfur removal efficiency, showing a better performance than the

commercial sorbents such as zeolites and activated carbon.

In 2000, a study of sulfur adsorption capacity of different sorbents is conducted

by Siemens Westinghouse Power Corporation (SWPC) in order to evaluate the

performance for removing sulfur components (Crespo, 2008). The sulfur adsorption

capacity is based on dimethyl sulfide (DMS) breakthrough profiles in a packed bed, in

which DMS was found to be the most difficult sulfur compound to be removed from the

natural gas. At high gas hourly space velocities of 60000/h, SulfatrapTM showed the best

performance out of all sorbents giving a sulfur adsorption capacity of 3.1 wt% at 720

min. Besides that, the saturation capacity which is defined as the total sulfur loading of

the sorbent is determined to be 3.9 wt%.

20

SYNTHANOL SDN. BHD.


Another critical feature of SulfatrapTM indicated from the analytical results is

that the sorbent does no catalyze any side reactions to form high molecular weight

sulfur compounds. According to Alptekin (2008), it showed that Zeolite-13X sorbent is

most likely to experience competition for adsorption sites and reduction of the available

sites for sulfur adsorption. As for SulfatrapTM-R2A, no competition for adsorption sites

from aromatic compounds is observed. Thus, this proved that SulfatrapTM-R2A is highly

selective to sulfur compounds only.

Most importantly, SulfatrapTM sorbent is totally regenerable by simply heating

up the sorbent bed up in the range of 300 – 425℃. On top of that, the sorbents still

managed to maintain a stable sulfur adsorption capacity after 10 – 31 cycles provided

the regenerations were carried out at 350 – 425℃ (Alptekin, 2008). Reuse and

regeneration of sorbents will able to reduce the operating cost significantly despite the

constant replacement of sorbents. Furthermore, it reduces the waste generation and

reduces the needs for landfill disposal. Based on Pierre (2008), TDA’s sorbents

replacement interval is approximately 3 years. Moreover, the required operating

temperature for SulfatrapTM is at ambient temperature which offers a great deal of

simplicity as compared to technologies which involve elevated temperature.

SulfatrapTM is a low cost, high sulfur removal capacity and regenerable sorbent

for removing sulfur component from natural gas at ambient temperature. It has low

affinity to hydrocarbons, does not alter the composition of the natural gas and sulfur

compounds that adsorbed on the sorbent. Hence, the sulfur compounds that remove by

adsorption are able to be recovered by Claus process. SulfatrapTM sorbent is non-

pyrophoric substance (a substance that will ignite spontaneously in air) and does not

contain toxic ingredients. In environmental and safety point of view, it does not require

any special handling for disposal and storage (Alptekin, 2006). However, the

composition of SulfatrapTM and mechanism for desulfurization of this technology is

unknown.

(vi) Activated Carbon Sorbents

Activated carbon is another sorbent applied commercially in removing H2S from

any gas streams (Armstrong, 2003). However, the key mechanisms involving the

21

SYNTHANOL SDN. BHD.


removal of H2S using activated carbon is not well-studied. Besides that, there is lack of

information on the critical features of activated carbon catalyst (Armstrong, 2003).

Activated carbon-based sorbent was also reported to have lower pre-breakthrough

capacity compared to zeolite-based sorbent and other commercially used sorbents

(Alphtekin, 2006).

In summary, Table 2.5 shows the comparison of each solid sorbent available in

the market.

Table 2.5: Comparison of solid sorbents

Parameters Zinc Oxide Zeolite SulfatrapTM [d] SulfaTreatTM

Unit PriceUDS 5 – 10

/ kg [a]

USD 1.8 – 3.25 / kg [a]

USD 4.54 – 11.34 / kg

USD 0.31/ kg [e]

Lifespan - - Regenerable Regenerable

Operating parameters:i) Temperatureii) Pressure

350 – 550℃ [b]

Room temperature (25

℃) [b]

30 kPa [b]

Room temperature (20

℃)34.47 kPa

> 177 ℃ [e]

3447 kPa [e]

Regeneration operating parametersi) Temperature

-350℃ with air

[b] 300℃ -

Performance:Sulfur adsorption capacity

1.2 wt% [c] 0.36 wt% [d] 3.1 wt% 12 wt%[f]

[a] (Foral et al,1993)[b] (Crespo et al, 2008)[c] (Copeland et al, 1998)[d] (Alptekin et al, 2006)[e] (SulfaTreat, 2011)[f] (Mi Swaco, 2002)

22

SYNTHANOL SDN. BHD.


2.1.3 Biological Process

Removal of H2S from sour gas via biological means involves the usage of an

adapted mixed microbial culture (consortium) which is capable of oxidizing the sulfide

species in sour gas (Srivastava et al., 2002). Many reported on the investigation of

utilizing chemoautotrophic bacteria which belongs to Genus Thiobacillus to remove

H2S (Srivastava et al., 2002). This technology is suitable for small scale operations in

which sulfur will be produced in a rate of 0.2 – 2 TPD with maximum 4 ppm H 2S

contained in the treated sweet gas (Srivastava et al., 2002). However, this is a novel H2S

removal technology and has not been applied in industrial scale (Srivastava et al.,

2002). Further research work has to be done before this technology can fit into the

current oil and gas industry without much limitation such as low capacity, high capital

and operating cost, and environmental issue while dealing with the microorganism

disposal.

2.1.4 Selection of Technology

According to Foral et al (1993), conventional chemical absorption / physical

solvents (liquid absorption process) are not economical for low H2S concentrations.

This is due to the fact that this technology is not suitable to be employed for gas streams

treatment with lower than 15% H2S in the feed stream (Nagl, 2007). Hence, solid bed

adsorption is more suitable to eliminate low concentration H2S in natural gas. A study of

comparison of H2S scavenging technologies is established in which a relative screening

index (SI) was developed considering investment and operating costs, and subjective

weightings of process reliability, ease of operation, operator acceptance, ease of spent

material disposal, and winterization requirements. This study showed that SulfaTreatTM

possessed the best rating among all the categories aforementioned with the lowest total

plant investment as shown in Figure 2.9 and Figure 2.10.

From Figure 2.9, SulfatreatTM had the highest score as compared to others, for

example zinc oxide and iron sponge. Although SulfatreatTM is a newly developed

technology, the process reliability is the highest amongst all. Furthermore, the ease of

operation is rather simple with the lead/lag configuration which offers greater utilization

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SYNTHANOL SDN. BHD.


Process reliability (PR) Winterization (W) Ease of operation (EOP) Operator Acceptance (OA) Ease of disposal of spent material (DOSM)

of the sorbents and flexibility in the scheduling of charging and removal of spent beds.

SulfatreatTM is classified as Class I non-hazardous material which could be landfill

directly. A potential application of the spent sorbents as a soil additive has been

proposed. Moreover, operating cost for SulfatreatTM eventually is the lowest as shown in

Figure 2.10.

The main advantage of this selected process is that the consumption of

SulfatreatTM is eventually dependent on the amount of H2S passes through the sorbents

bed. Ability to adapt changes in alteration of operating parameters or preferences

without the need of additional capital requirement and system modification is another

advantage of SulfatreatTM sorbents. In short, the advantages of this technology are such

as long bed life or life span of SulfatreatTM, predictable pressure drops, consistent

product performance, environmental friendly, safe handling and simple operation (Mi

Swaco, 2010).

24

Figure 2.9: Comparison of Sulfatreat with other commercialized sorbents (Foral et al, 1993).

SYNTHANOL SDN. BHD.


Figure 2.10: Comparison of operating cost for sulfur scavenging process (Foral et al,

1993).

2.2 Syngas Production Technology

Module (M), which is defined by the stoichiometric ratio (H2 – CO) / (CO +

CO2), is the parameter used to characterize the synthesis gas. A module of 2 defines a

stoichiometric synthesis gas for the formation of methanol (Petersen et al., 2008). A

module below 2 should be avoided because it will result in the formation of byproducts

and also a loss of synthesis gas as increased purge (Hansen and Nielsen, 2008). Besides

module, H2O to C ratio, CO to CO2 ratio and concentration of inerts are some of the

important properties for the production of synthesis gas. If the CO to CO2 is very high,

the rate of reaction and thus the achievable per pass conversion will increase. By this

way, this reduces the formation of water and the rate of deactivation of the catalyst in

the pre-reformer, steam reformer and autothermal reformer will decrease (Arthur,

2010).

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SYNTHANOL SDN. BHD.


On the other hand, if the concentration of inerts such as methane, ethane,

nitrogen and argon is very high in the synthesis gas, it will greatly affect the partial

pressure of active reactants resulting in a decrease in the rate of reaction. Therefore, the

ideal synthesis gas will contain a low content of inerts and a high CO to CO2 ratio. Due

to the high H2O to C ratio, the syngas produced contains a large amount of H2 content in

the conventional reformer leading to a high module number which is not suitable for

methanol production (Aresta, 2003).

Synthesis gas (syngas) is basically a mixture of hydrogen (H2), carbon monoxide

(CO) and carbon dioxide (CO2). Synthesis gas is produced from a number of different

feedstocks such as natural gas, coal, biomass, naptha and heavy residuals (Arthur,

2010). However, among all feedstocks, the most applicable in the methanol production

is natural gas. A number of different technologies are currently available for the

production of synthesis gas and also have been described in detail in most of the

literatures. For instance, pre-reforming, conventional steam reforming (one step

reforming with fired tubular reforming), autothermal reforming (ATR), combined

reforming (two step reforming), gas heated reforming, heat exchange reforming and so

forth.

2.2.1 Adiabatic Pre-reformer (APR)

Adiabatic pre-reforming is a process used for the reforming of feedstock which

ranges from natural gas to heavy naphtha (Logdberg and Jakobsen, 2010). It is a key

element in an optimised design of the synthesis gas generation unit in a gas-to-liquid

plant (Petersen et al., 2004). A feedstock that is rich in higher hydrocarbons first needs

to be treated in a pre-reforming step. This is to convert the heavy hydrocarbons in the

feed into methane, hydrogen and carbon oxides (Ijaz, 2008). In addition, water gas shift

and methanation reactions will occur simultaneously. Some methane might be steam

reformed in this process as well. The extent of reforming depends on various factors,

namely the feed preheat temperature, operating pressure, feed gas composition and

steam to carbon ratio (Ijaz, 2008).

The reactions which occur in this step include:

CO+3H 2⇌CH 4+ H 2 (Methanation reaction)

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SYNTHANOL SDN. BHD.


CO+3H 2O⇌ CO2+H 2(Water−gas shift reacti on)

Cn Hm+( m2−1) H 2O → nCO2+(m−1) H 2(Hydrocrackingreaction)

Natural gas is first fed to this process after desulphurisation and preheating to

its desired reactor inlet temperature. Subsequently, the effluent from the pre-reforming

step is further preheated and fed to a downstream reformer. A pre-reformer is typically

operated adiabatically at temperatures between 320 and 550 °C whereby a heat

exchanger coil that is installed in the convection duct of the steam methane reformer

may be advantageously used for preheating purposes (Ijaz, 2008).The heat content of

the feed stream will be utilized to drive the steam reforming reaction at low

temperatures (Arthur, 2010).

The operation of a pre-reformer within its allowable temperature range is

important due to the formation of a whisker type carbon which will occur above the

upper temperature limit. On the other hand, operation below the lower temperature limit

may result either in a polymeric type of carbon formation (gum) or lack of sufficient

catalyst activity (Petersen et al., 2004). The operating pressure, however, ranges from 3-

4 MPa with a steam to carbon ratio of 0.5 to 3.5 (Ijaz, 2008).

For heavy feedstock such as naphtha, the overall prereforming process is often

exothermic whereas lighter feedstock such as LPG and natural gas may result in an

endothermic, thermoneutral, or exothermic reaction (Petersen et al., 2004). This may

lead to a lead to a net temperature drop depending on the content of higher

hydrocarbons (Ijaz, 2008).

The pre-reforming step has several advantages. The removal of the higher

hydrocarbons from natural gas enables a higher feed temperature to further reforming

processes without having to face the risk of thermal cracking in the preheater coil. A

higher feed temperature entering subsequent down-stream reformers reduces the oxygen

consumption and carbon efficiency (Petersen et al., 2004). Other than that, the

production capacity of the plant may be increased because by installing a new pre-heat

coil between the pre-reformer and the steam reformer the load on the reformer is

reduced. This may be used as a capacity increase or with unchanged capacity, result in a

decrease in firing (Ijaz, 2008). Besides that, the chemisorption of sulfur to the Ni-

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SYNTHANOL SDN. BHD.


catalyst would be favourable due to the fact that temperature in the pre-reformer is

relatively low. Therefore, traces of sulfur from the desulfurization unit will be trapped

in the pre-reformer. This can increases the life-time of the tubular steam reforming

catalyst since there would not be any sulphur poisoning at the top layer of the catalyst

(Logdberg and Jakobsen, 2010).

Carbon formation from higher hydrocarbons is an irreversible reaction. It can

only take place in the first part of the reactor where there is the highest concentration of

C2+ compounds. Also, the risk of carbon formation is most prone in the reaction zone

where the temperature is the highest (Petersen et al., 2004). Other than that, the

conversion of higher hydrocarbons to methane is crucial as they tend to become more

reactive in the steam reforming process. This would lead to carbon formation and thus

to deactivation of the catalyst employed (Ijaz, 2008). In order to limit the carbon

formation, the ratio of steam to higher hydrocarbons can be reduced and temperature

increased (Petersen et al., 2004).

2.2.2 Steam Methane Reforming (SMR)

Process Description

The dominating technology for the production of syngas from a methane

feedstock is the reaction with steam at high temperatures. The conventional term for this

method is called steam methane reforming (Ijaz, 2008). Here, the feedstock is

catalytically cracked in the absence of oxygen with the addition of water and possibly

carbon dioxide (Hansen and Nielsel, 2008). Typical feedstock for this process ranges

from natural gas and LPG to liquid fuels including naphtha (Petersen et al., 2004).

When natural gas is subjected to steam reforming, it tends to form a mixture of

hydrogen and carbon oxides which is crucial in the subsequent stages of methanol

production (Cheng and Kung, 1994). Two principal reactions that take place in the

steam reformer include:

CH 4+H 2 O⇌ CO+3H 2(Reforming reaction)

CO2+H 2O⇌ CO2+H 2(Water−gas shift reaction)

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SYNTHANOL SDN. BHD.


The predominant reforming reaction is strongly endothermic whereas the

accompanying water-gas shift reaction is moderately exothermic (Ijaz, 2008).

Therefore, the overall steam reforming process is highly endothermic and is carried out

at high temperatures ranging from 800 ºC - 900 ºC and at pressures between 15 and 36

bar over a Ni/Al2O3 catalyst (Logdberg and Jakobsen, 2010). The product gas leaving

the reformer at an elevated temperature can then be cooled in a process gas waste heat

boiler to produce process steam for the reformer (Ijaz, 2008).

Although steam reforming is valid as a stand-alone process (Petersen et al.,

2004), it by itself is not the preferred technology for production of synthesis gas for

large-scale GTL applications. It is commonly used in combinations of various oxygen

or air-blown partial oxidation processes (Petersen et al., 2004). This is because large-

scale steam reformers have a poor economy of scale as compared to processes based on

partial oxidation and air separation as they require large heat input (Petersen et al.,

2004).

Other than that, the syngas produced via conventional steam reforming

typically has a stoichiometry number, SN of between 2.6 and 2.9. However, for

methanol production, the preferred SN value for the produced syngas is 2. One of the

methods used to lower this value is by the addition of carbon dioxide or by combined

reforming (Section 2.2.4) (Ijaz, 2008).When the feed is natural gas without carbon

dioxide addition, the SN is close to 3which is far from the desired value of 2. With

carbon dioxide addition, lower values of SN can be obtained with a lower energy

consumption of about 5 – 10% as compared to a conventional plant (Hansen and

Nieisel, 2008)

The steam to methane ratio (S:M) is another important parameter to be closely

monitored in the steam reforming process. Figure 2.11 shows that a high S:M ratio in

the feed is required to give high conversions especially at elevated pressures (Petersen

et al., 2004). If this ratio is too low, carbon deposits will occur and this will

subsequently deactivate the catalyst by coking. Large carbon deposits may also block

the tubes and cause hot-spots. A common steam/carbon ratio lies between 2.5 and 4.5. A

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SYNTHANOL SDN. BHD.


higher ratio helps shift the reforming equilibrium towards the products, hence

increasing the methane conversion.

Figure 2.11: Relationship among steam reforming temperatures, S:M ratios and

methanol conversion (Petersen et al., 2004).

Equipment Description

In industrial practice, steam reforming is mainly carried out in reactors referred

to as steam reformers (Petersen et al., 2004) which are essentially large process furnaces

in which catalyst-filled tubes are heated externally by direct firing to provide the

necessary heat for the reactions taking place inside the reformer tubes (Cheng and

Kung, 1994).

A conventional steam reformer consists of two sections – a convection and a

radiant section. The reforming reaction of the process gas takes place in the radiant

section which contains several rows of vertical tubes. Steam is mixed with the process

gas prior to entering these tubes. Here, the process gas is gradually heated to about

800ºC via heat exchange with the hot flue gas in the firebox (Logdberg and Jakobsen,

2010).

However, only 50% of the heat produced by the combustion in the burners is

transferred to the process gas. This heat is needed to drive the reaction and to bring the

products to the exit temperature. The other 50 % of heat liberated exits the system in the

hot flue gases from the burners. This remaining unabsorbed heat in the reforming

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SYNTHANOL SDN. BHD.


section must be recovered in the convection section of the furnace to ensure a

thermodynamically efficient operation. The overall furnace efficiency can be as high as

92.93% whereby the flue gases are released at about 150°C (Cheng and Kung, 1994).

After the flue gas has supplied its heat to all the reactor tubes, it passes through

the convection section to be further cooled by heating other streams such as feed to

other processes, combustion air, boiler feed water as well as for steam production (Ijaz,

2008). Since most upstream and downstream processes obtain heat input (preheating)

from the hot flue gas in the convection section, the tubular reformer is also commonly

seen as an energy converter (Petersen et al., 2004). The fuel used for combustion in the

firebox is usually the same hydrocarbon as the process stream, namely natural gas.

If the production of surplus energy is unnecessary, smaller tubes can be installed

inside the existing reformer tubes. The catalyst is placed in the space between the two

tubes where the combined stream of steam and natural gas enters (Logdberg and

Jakobsen, 2010). At the end of the reformer tube, the gas enters the smaller tubes and

transfers some heat to the catalysts before exiting at the top. By implementing this, the

number of tubes as well as the total surface area can be reduced by approximately 20%


Equipment Design

As history goes, until the 1980s, most reformer furnaces were constructed using

centrifugally cast 25% chromium and 20% nickel (HK-40) alloy tubes. However, a

higher strength 25% chromium and 35% nickel-niobium (HP modified) cast tube has

been intensively used in recent years as it is found to be stronger with improved stress-

to-rupture properties, thus resulting in thinner tubes containing less net metal for the

same design tube life (Cheng and Kung, 1994).

Steam reformers can be said to be ‘heat flux limited’ due to the fact that the

reactor is usually limited by heat transfer considerations and not by reaction kinetics.

The number of tubes and their dimensions are designed to achieve the desired heat flux

profile whereby the amount of catalyst should be sufficient to achieve the desired level

of conversion (Van Den Oosterkamp and Van Den Brink, 2010).

31

SYNTHANOL SDN. BHD.


In practice, a SMR unit may contain from 40 up to 1000 tubes, each typically 6-

12 m long with inner diameters of 70-160 mm (Ijaz, 2008).The wall thickness of the

tubes is between 10 – 20 mm. The small tube diameters are crucial in order to achieve

the highest possible heat flux to the catalyst and thus, achieve the highest possible

capacity for a given amount of catalyst (Logdberg and Jakobsen, 2010).

A well-designed reformer with good heat transfer characteristics would still

experience high heat fluxes resulting in a significant film temperature drop between the

inside reformer wall temperature and the bulk gas temperature. Therefore, it is

necessary to evaluate coking tendencies at the reformer at wall temperature conversion

(Van Den Oosterkamp and Van Den Brink, 2010).

Figure 2.12: Different burner configurations used in steam reformers (Logdberg and

Jakobsen, 2010).

The four types of burner configurations used in steam reformers include top

fired, bottom fired, terrace wall and side fired burners. The graphical interpretations of

these burners are as shown in Figure 2.12. The burner geometry, flame length and

diameter, tube-to-tube and row-to-row spacing, fired tube length and distance from the

flame to the reformer wall determines the homogeneity of the heat transfer to the tubes

(Logdberg and Jakobsen, 2010). Therefore, the selection of the type of burner

configuration is extremely important in terms of heat flux and hence, capital investment

conversion (Van Den Oosterkamp and Van Den Brink, 2010).

The following would include descriptions of each burner type:

32

SYNTHANOL SDN. BHD.


The bottom fired type is today considered out-dated as it was only widely used in

the past. It gives an almost constant heat flux profile along the length of the tube. A

substantial margin is required on the tube design temperature in order to limit the

outlet temperature since the tubes are hot at the bottom. This type provides an easy

access to the burners (Petersen et al., 2004).

A modification of the bottom fired type resulted in the terrace wall reformer. This

type was found to have slightly lower tube wall temperatures. However, problems

can arise at the 'pinch point' in the middle of the furnace. This is due to the fact that

the tubes are subject to both radiations from the burners and to enhanced convection

from the flue gas at this point (Petersen et al., 2004).

A more widely used burner type would be the top fired reformer. Top-fired

reformers have several parallel rows of tubes (Logdberg and Jakobsen, 2010) with

burners mounted in the furnace ceiling between the tubes as well as between the

tubes and the furnace wall (Petersen et al., 2004). The tubes are heated via radiation

from the flames and the hot flue gas and by convection (Logdberg and Jakobsen,

2010). In some designs, the feed gas and hot flue gas flow in parallel down the

length of the tube. The manifolded tubes collect the synthesis gas, which passes

back up through the furnace in riser pipes. This is done in order to collect more heat

before passing into the effluent transfer line and out of the reformer (Cheng and

Kung, 1994). Other top fired designs allow a bottom exit where gas exits the

catalyst filled tubes through pigtails before passing to external collection manifolds.

The flue gas is pulled out through the convection section whereby additional heat is

extracted to increase the overall furnace efficiency before final discharge to the

atmosphere (Cheng and Kung, 1994). The top fired reformer has the highest heat

flux where the temperature of metal is at its maximum. As the catalyst deactivates, a

slight increase in temperature in the lower end of the tube makes it possible to retain

the productivity. However, this will result in a large temperature increase in the top

of the tube. Therefore, top fired reformers must be designed with a considerable

margin above the maximum temperature at the start of the run (Logdberg and

Jakobsen, 2010).

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SYNTHANOL SDN. BHD.


The side fired reformer can only have one row of tubes and heat transfer is mainly

by the radiant side-wall. The side-fired reformer not only allows for a better

temperature control but also has the maximum temperature at the outlet of the tube.

The side-fired reformer has a higher average heat flux than the top fired and the

highest heat flux occurs at a rather low temperature. Other than that, this reformer

gives very low emissions of NOx in the flue gases due to the short residence time in

the flames. Moreover, a decrease in catalyst activity will lead to an increase in

temperature in the upper part but the temperature will still be highest in the lower

end. Therefore, the reformer does not have to be designed for much higher

temperatures than at the start of the run (Logdberg and Jakobsen, 2010).

Catalyst Details

As mentioned before, the reactor tubes in the steam reformer contain nickel-

based catalyst (Ijaz, 2008). Since methane is a very thermodynamically stable molecule

even at high temperatures, the catalyst is needed to reduce the operating temperature

and hence, decrease the tube stresses resulting from high pressure and high

temperatures. The methane reforming is a first-order reaction irrespective of pressure.

At high temperatures, the overall rate can be limited by pore diffusion. However, at low

temperatures, the molecular diffusion rate is much higher than the reaction rate so that

the catalyst activity can be fully used. At high temperatures, the overall rate in steam

reforming is limited by the heat transfer (Logdberg and Jakobsen, 2010). The Ni-

catalyst commonly used is in the form of thick-walled Raschig rings with dimensions 16

mm in diameter and height, and a 6 – 8 mm hole in the middle. The limits of such

catalysts will be reached if the heat load per unit area is too high. Subsequently, smaller

particles will be necessary in order to make use of more of the catalyst. However,

smaller particles will result in an increased pressure drop. Therefore, special packing

shapes such as spoked wheels or rings with several holes will have to be used


34

SYNTHANOL SDN. BHD.


2.2.3 Autothermal Reforming (ATR)

Process Descriptions

Autothermal reforming is the reforming of light hydrocarbons in a mixture of

steam and oxygen in the presence of a catalyst. ATR requires O2 which is produced

from an air separation unit (ASU). A lower H2 to CO ratio would then be obtained by

the addition of O2. Owing to high investment costs for the separation of the oxygen

from air, the autothermal reformer is usually not standalone. It is normally located

downstream a steam reformer acting as a secondary reformer in order to further reform

the unreacted methane from the primary reformer to achieve a stoichiometric ratio of

synthesis gas (Logdberg and Jakobsen, 2010).

Nearly pure oxygen (99.5%) is injected rather than air because the presence of

excessive N2 as an inert in the syngas would overburden the compressors in the latter

stage and hence retard methanol synthesis leading to a low overall efficiency (Cheng

and Kung, 1994; Petersen et al., 2004). By introducing O2 into the ATR, excess H2 is

combusted resulting in a drop of stoichiometric ratio from 3.0 to 1.8 which is much

nearer to the desired value of 2.0 (Logdberg and Jakobsen, 2010).

In the autothermal process for syngas production, the heat of reaction is supplied

by partial oxidation of natural gas for subsequent endothermic reforming reaction. The

overall process is known as autothermal. Autothermal reforming is a low investment

process using a simple reactor design (Haid and Koss, 2001). No tubular steam reformer

is required unlike the conventional steam reforming. Typical process conditions are 950

– 1100 oC and 20 – 40 bar (Haid and Koss, 2001; Logdberg and Jakobsen, 2010).

Besides that, the steam to carbon ratio, which is based on the total feed, is found

to be in the range 2.0 to 2.5 (Petersen et al., 2004). Low steam to carbon ratio will result

in an increase of CH4 leakage (unconverted methane in the effluent of ATR) in the

synthesis gas. On the other hand, oxygen to carbon ratio is between 0.6 and 1.5

(Logdberg and Jakobsen, 2010). The synthesis gas produced by autothermal reforming,

which is rich in carbon monoxide and 15 – 20% deficient in hydrogen, has a

stoichiometric ratio of 1.7 to 1.8 (Hansen and Nielsen, 2008; Petersen et al., 2008). To

adjust the module to a value of 2.0, there are a few adjustments which could be

35

SYNTHANOL SDN. BHD.


performed. For instance, it can be done by removing CO2 from the synthesis gas,

recovering hydrogen from the purge gas by membranes or a pressure swing adsorption

unit (PSA) or recycling the recovered hydrogen from synthesis gas (Petersen et al.,

2004; Hansen and Nielsen, 2008; Petersen et al., 2008). Besides all these methods, the

amount of oxygen entering the ATR could be adjusted to adjust the syngas so that a

module of 2.0 is achieved (Hansen and Nielsen, 2008).

The overall chemical reactions involved in the whole ATR reactor are shown in

the following equations.

Combustion zone:

CH 4+12

O2⇌CO+2 H 2 ∆ H ro=−35.67 kJ /mol

2 H 2+O2⇌ 2 H2 O ∆ H ro=−483.66 kJ /mol

Catalytic zone:

CH 4+H 2 O⇌CO+3 H 2 ∆ H ro=206.16 kJ /mol

CO+ H 2O⇌CO2+H 2 ∆ H ro=−41.15 kJ /mol

Figure 2.13: Autothermal Reformer (Logdberg and Jakobsen, 2010).

36

SYNTHANOL SDN. BHD.


Equipment Descriptions and Design

The ATR reactor consists of a refractory-lined pressure vessel. As the name

implies, it can stand higher pressures and temperatures than the steam reformer. The

reactor vessel is lined on the inside with refractory which insulates the steel wall of the

pressure vessel from high temperature reaction environment (Petersen et al., 2004). The

refractory consists of several layers with different materials and insulation materials.

Nowadays, a refractory design with three layers of refractory is used to further protect

the reactor from any cracks in the refractory layers.

Basically, the reactor space comprises three different zones such as a burner, a

combustion chamber and a fixed catalyst bed in which different reactions occur as

shown in Figure 2.13 (Petersen et al., 2008). The gas flows from the top to the bottom

through a catalyst bed supported by a ceramic arch (Uhde, 2006). Firstly, the burner

provides good mixing of the feed streams and the oxidant in a turbulent diffusion flame.

The core of the flame has a very high temperature which can reach more than 1000 oC.

Effective mixing at the burner nozzles and also recirculation of the reacted gas from the

thermal zone to the burner can protect the refractory and burner from the hot flame core

and gases from the combustion zone (Petersen et al., 2004; Logdberg and Jakobsen,

2010). With the use of oxygen or enriched air as oxidant, the speed of flame will be

much faster than that for air flames. As a proof, the position of the oxygen flame is

closer to the nozzles of burner as compared to the air flame (Petersen et al., 2004). The

residence time in the burner is typically short (1 – 3 seconds) (Van Den Oosterkamp and

Van Den Brink, 2010).

Next, in the combustion zone, the natural gas reacts with oxygen/steam by sub-

stoichiometric combustion in a turbulent diffusion flame as shown in the equation 8.

The combustion conditions are sub-stoichiometric since the overall oxygen to

hydrocarbon ratios vary between 0.6 and 1.5 (Logdberg and Jakobsen, 2010). H2 formed

from equation 8 will be burnt to water according to equation 9. The gas exiting the

combustion chamber in the ATR contains a considerable amount of methane and other

gas components (Petersen et al., 2004). It is ensured that the gas and temperature

distribution must be homogeneous before entering the catalyst bed in catalytic zone

37

SYNTHANOL SDN. BHD.


(Petersen et al., 2004). Inhomogeneity of gas will cause a greater distance to

equilibrium and hence the concentration of methane in the outlet gas is increased.

Lastly, the catalytic zone is a fixed bed in which the hydrocarbons are finally

converted through heterogeneous catalytic reactions including steam methane reforming

and water gas shift reaction (Logdberg and Jakobsen, 2010). A layer of protecting tiles

is usually placed on top of the catalyst bed to protect it from the very intense turbulent

flow in the combustion chamber. The catalyst bed is operated in the range of 950 – 1400 oC. According to Pina and Borio (2006), the reported temperature value from industry in

the catalytic zone was found to be 950 oC.

Most reforming catalysts are based on nickel as the active material (Petersen et

al., 2004). Besides nickel, Cobalt, Ruthenium, Rhodium and noble metals are able to

catalyse the reforming reactions as well. However, they are generally very expensive to

be used industrially although they have higher activity per unit metal area than the

conventional nickel catalysts (Petersen et al., 2004; Nielsen, 2008). Thus, the common

catalyst used in the catalytic zone is nickel supported on an alumina base due to high

thermal resistance, high thermal stability and not prone to deactivation (Petersen et al.,

2004; Nielsen, 2008). Therefore, sufficient strength could be achieved at the high

operating temperatures Petersen et al., 2004). However, seeing as the catalyst is exposed

to high operating temperatures, the nickel metal is subjected to a high degree of

sintering (Petersen et al., 2004). The catalysts used in the catalytic zone should be

optimised in order to maximise the heat transfer and strength at a low pressure drop

(Petersen et al., 2004). The shape and size of the catalyst particles should be optimised

as well to achieve maximum activity with a minimum pressure drop. This causes a

compromise between low particle diameter and high void fraction. According to Nielsen

(2008), the optimum is a catalyst bed of particles with large diameter and with high void

fraction.

The catalyst bed brings the steam methane reforming and water gas shift

reactions to equilibrium over the catalyst bed in the synthesis gas and destroys soot

precursors (Petersen et al., 2008). Therefore, the operation of ATR is soot-free. Also,

soot-free operation could be achieved through the optimised burner design. Formation

of soot precursors such as poly-aromatic hydrocarbons (PAH) would greatly decrease

38

SYNTHANOL SDN. BHD.


the carbon efficiency of the methanol process (Petersen et al., 2004). Besides reducing

the soot formation, excessive temperatures could be avoided with a careful design of

burner and combustion chamber (Petersen et al., 2004).

In addition to that, the syngas is completely free of oxygen (Logdberg and

Jakobsen, 2010). It is found that the overall reaction rate is controlled by the transport

rate of reactants through the gas film surrounding the catalyst pellets (Logdberg and

Jakobsen, 2010). Since the catalytic reaction is extremely fast, the process is carried out

at high space velocity. Higher space velocity will directly reduce the gas film thickness

surrounding the catalyst pellets resulting in a better heat and mass transfer with the

catalysts. The size and shape of catalyst particle is optimised to achieve high activity per

unit area, high selectivity and low pressure drop in order to reduce any side reactions.

By far and large, the ATR or secondary reformer is operated close to adiabatic

condition and thus the temperature is determined from the adiabatic energy balance

(Logdberg and Jakobsen, 2010). For the design of ATR (combustion and catalytic

zones), it is crucial to reduce the hot spots on the pressure shell (reactor vessel) which

otherwise could result in a much higher rate of creep rupture and catalyst sintering or

plugging (Van Den Oosterkamp and Van Den Brink, 2010).

There are several advantages of using this technology. As compared to

conventional steam reforming, autothermal reforming achieves a reduction of 30% and

80% in CO2 and NOx emissions respectively (Haid and Koss, 2001). Besides that, the

thermal efficiency (ratio of lower heating value of reformed gas to that of the

hydrocarbon feed) is higher (88.5%) than that of conventional steam reforming (81%)

and also than that of partial oxidation (83.5%) (Logdberg and Jakobsen, 2010).

Unlike steam reforming, the maximum temperature is not limited by the tube

material but it is limited by the stability of the catalyst and also refractory lining of the

reactor (Logdberg and Jakobsen, 2010). Furthermore, autothermal reforming is more

flexible than tubular reforming since it can operate at a higher temperature to

compensate for the increase in methane slip (unconverted methane from the primary

reformer).

39

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By installing ATR in the downstream of SMR, the heat load of steam reformer

could be substantially reduced approximately 70%. As a result, a smaller primary

reformer and less fuel would be required. Therefore, this indirectly reduces the size of

the related equipment in the flue gas duct area in the convection side of steam reformer

(Uhde, 2006).

2.2.4 Combined Reforming

Combined reforming is usually applied for heavy natural gases and oil-

associated gases (Lurgi, 2006). Heavy natural gas consists of higher hydrocarbons such

as ethane and propane besides just methane. The required stoichiometric number cannot

be obtained by pure autothermal reforming only. The two step reforming process

(combined reforming) features a combination of steam reforming (primary reforming)

followed by autothermal reforming (secondary reforming) with oxygen providing the

heat source (Uhde, 2006). The basic objective of combining these two reforming

technologies is to adjust the stoichiometric ratio of synthesis gas to obtain the most

suitable composition (a module of 2 for methanol synthesis).

The remainder of the feed gas from the desulphuriser is mixed with the steam

reformed effluent (from the primary reformer) in the autothermal reformer. Secondary

reforming is a process in which partially converted process gas from a tubular steam

reformer is further converted by means of internal combustion (Logdberg and Jakobsen,

2010). Combustion in the upper zone of the secondary reformer increases the

temperature of the partially combusted gas. The temperature of the combusted gas will

then decrease rapidly in the catalytic zone whereby the endothermic process absorbs

heat as it progresses axially along the catalyst bed (Cheng and Kung, 1994). From here,

the main advantage of the combined reforming is the original feed gas bypass of the

steam reformer (Lurgi, 2006). By bypassing some of the reforming duty from the

primary reformer to the secondary reformer, the size of primary reformer and fired duty

are greatly reduced (Cheng and Kung, 1994). The similar descriptions of steam

reforming (primary reforming) and autothermal reforming (secondary reforming) are

described in Sections 2.2.2 and 2.2.3.

40

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2.2.5 Heat Exchange Reforming

Heat exchange reforming applies a concept whereby the process gas supplies

part of the heat required to the tubes via heat exchange. When two reformers are

combined, the heat needed in the tubular steam reformer is obtained from the hot

product gas from the secondary reformer. This concept can be used for production of

hydrogen or syngas for the methanol synthesis (Logdberg and Jakobsen, 2010).

This heat is needed for the endothermic steam-reforming process and is

delivered by convective heat transfer from hot syngas product and flue gas conversion

(Van Den Oosterkamp and Van Den Brink, 2010). This method of reforming eliminates

the expensive fired reformer (Logdberg and Jakobsen, 2010).

However, only medium pressure steam can be recovered from the syngas plant

and electricity for the syngas compressor must be imported (Logdberg and Jakobsen,

2010). Plants that use this concept produce much less steam to be exported because

much more heat integration takes place in the reactor itself (Van Den Oosterkamp and

Van Den Brink, 2010). A significant number of possible combinations exist when it

comes to heat exchange reformers. These reformers which are heated by process gas are

always installed in combination with other reformers, namely a fired tubular reformer or

an air or O2-blown auto-thermal reformer (Petersen et al., 2004). Over the years,

several reactor concepts which make use of this convective heat transfer concept have

been developed.

The Gas-Heated Reformer (GHR) concept uses the heat content present in the

synthesis gas, which is being produced by an ATR. This reactor typically consists of a

number of catalyst-filled tubes, each with a central bayonet tube. The annular space

between these concentric tubes is filled with catalyst. The feed gas enters the top of the

reactor vessel and flows through the catalyst-filled annular space and then back through

the central tube while simultaneously giving off heat to the incoming feed gas

conversion (Van Den Oosterkamp and Van Den Brink, 2010). The gas then passes on to

the ATR or secondary reformer. In order to increase the heat transfer coefficient the

outside surface of the outer tube would be designed as a finned surface conversion (Van

Den Oosterkamp and Van Den Brink, 2010).

41

SYNTHANOL SDN. BHD.


A gas and steam mixture is fed to the catalyst tubes whereby the reaction takes

place. The ATR which is fired by oxygen or air then receives this partially reformed

gas. When the reforming reaction is completed, the resulting synthesis gas with high

heat content is passed to the shell side of the GHR. The synthesis gas then supplies the

heat required for the reforming reaction conversion (Van Den Oosterkamp and Van Den

Brink, 2010).

Another type of heat exchange reformer is the Convection Reformer. In this

reformer, flue gas which flow upwards on the outside of the tubes as well as the

reformer gas flowing upwards inside the tube would be the main sources of heat for the

reaction occurring (Logdberg and Jakobsen, 2010). The Topsoe convection reformer is

designed to have a single burner which is separated from the tube section. Since the

radiant tube section and the hot part of the convection section are combined in a

relatively small unit, it is termed as a convection reformer (Logdberg and Jakobsen,

2010). After heat exchange, the exit temperature from the reformer is approximately

600 ºC for both product gas and flue gas. This reduction in temperature signifies that 80

% of the fired duty is utilized in the process. This is much higher than the 50 %

achieved in a conventional steam reformer (Logdberg and Jakobsen, 2010).

A problem associated with heat exchange reforming would be the contact

between CO-rich gases with metals at high temperatures. This poses the risk of metal

dusting corrosion. The formation of carbon is possible via the exothermic Boudouard

reaction especially at temperatures below which the mixture satisfies the Boudouard

reaction equilibrium. A CO rich gas has a high Boudouard temperature and this makes it

easier for this reaction to be catalysed by hot metal surfaces (Logdberg and Jakobsen,

2010). Therefore, it is important that a metal surface of a slightly lower temperature

than a gas mixture does not come in contact with a gas mixture of high Boudouard

temperature. Carbon deposition on the metal would result in a big risk of metal

corrosion. Furthermore, if carbon is deposited on the catalyst, this will subsequently

lead to catalyst deactivation (Logdberg and Jakobsen, 2010).

42

SYNTHANOL SDN. BHD.


2.2.6 Partial Oxidation (POX)

Partial oxidation is often applied for gasification of heavy oil (Petersen et al.,

2004). However, all hydrocarbons are possible as feedstocks. Thus, this process is very

versatile which can convert a wide range of hydrocarbon feedstocks to synthesis gas.

The oxidant and the hydrocarbons are mixed in a reactor where the reactants are

allowed to react at very high temperatures in the range of 1300 – 1400 oC (Logdberg

and Jakobsen, 2010; Petersen et al., 2004). High exit temperatures from the gasifier will

minimise the formation of soot and also ensure the complete conversion of feedstocks

(Petersen et al., 2004). The operating pressure is found to be around 25 – 40 bar

(Petersen et al., 2004). The H2/CO ratio is lower as compared to conventional steam

reforming or autothermal reforming because no water is added in partial oxidation

process (Logdberg and Jakobsen, 2010). Partial oxidation of natural gas is usually used

in small plants and in regions where natural gas is cheap (Logdberg and Jakobsen,

2010).

Since the partial oxidation is a slightly exothermic reaction, the partial oxidation

reactor would be more energy efficient as compared to the energy intensive steam

reformer (Cheng and Kung, 1994). Besides that, seeing as the reaction proceeds fast, the

size of the reactor will be greatly reduced (Logdberg and Jakobsen, 2010). Partial

oxidation can be carried out with or without a catalyst. When a catalyst is used, the

reaction temperature will be lowered. The reaction will still achieve equilibrium since

the catalyst lowers the activation energies (Logdberg and Jakobsen, 2010). The resulting

gas is cooled by steam production and carbonaceous by-products such as soot are

discarded by washing. The carbonaceous by-products must be removed since they could

affect the carbon efficiency. In general, this process is widely used if the feedstock

contains a variety of components including the heavy oil. Table 2.6 summarises all the

current reforming technologies for the syngas production.

Table 2.6: Summary of Current Reforming Technologies for Syngas Production.

Reforming Technology

Operating Conditions

Advantages Disadvantages References

Adiabatic Reforming

(APR)

350 – 550 oC 30 – 40 bar Pressure drop

≤0.4 bar

Enables a higher feed temperature to further reforming

- Petersen et al. (2004); Logdberg and Jakobsen (2010)

43

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processes Traces of sulphur

will be trapped Reduces the

oxygen consumption

Steam Methane

Reforming (SMR)

800 – 900 oC 15 - 36 bar H2O/C: 2.5-4.5

Most extensive industrial experience

Oxygen is not required

Highest air emission (CO2

and NOx) High steam and

energy requirements

Cheng and Kung (1994); Petersen et al. (2004); Nielsen (2008); Logdberg and Jakobsen (2010)

Pure Autothermal Reforming

(ATR)

950 – 1100 oC 20 – 40 bar H2O/C: 2 – 2.5 O2/C: 0.6 – 1.5

Low methane slip Stoichiometric

ratio of syngas Lower process

temperature than POX

Oxygen is required

Limited industrial experience

Cheng and Kung (1994); Petersen et al. (2004); Pina and Borio (2006); Logdberg and Jakobsen (2010)

Two-step (Combined) reforming

Steam reformer: 800 – 900 oC 30 – 40 bar H2O/C: 2.5 – 3.0

ATR: 1000 – 1050 oC 20 – 40 bar 1 mol% CH4 slip O2/C: 0.6 – 1.5

Size of SMR is reduced

Steam reformer load is reduced

Overall feed and fuel consumption is lower than SMR

Stoichiometric ratio of syngas

Low methane slip Lower process

temperature than POX

Plant cost is 15% more than SMR

Higher process temperature than SMR

Increase the plant complexity

Oxygen is required

Lower CO2 and NOx emission than SMR

Cheng and Kung (1994); Petersen et al. (2004); Pina and Borio (2006); Uhde (2006); Hansen and Nielsen (2008); Logdberg and Jakobsen (2010)

Heat exchange reforming

600 oC Eliminates the expensive fired reformer

Heat integration takes place in the reactor itself

About 80% of fired duty is utilized in the process.

Only medium pressure steam can be recovered

Electricity for the syngas compressor must be imported

Contact between CO-rich gases with metals at high temperatures poses the risk of metal dusting

Van Den Oosterkamp and Van Den Brink (2010)

Logdberg and Jakobsen (2010)

44

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corrosion

Partial Oxidation

(POX)

1300 – 1400 oC 25 – 40 bar

Feedstock desulphurisation is not required

Low methane slip

Very high process temperature

Oxygen is required

Soot formation Increase the

process complexity

Petersen et al. (2004); Logdberg and Jakobsen (2010)

2.2.7 Economics, Safety and Environmental Considerations for Reforming Process

The investment in the syngas generation accounts for 50% – 60% of the total

investment in a methanol production plant (Hansen and Nielsen, 2008). Natural gas

reforming is the cheapest and most efficient syngas generation technology as compared

to other feedstocks such as coal gasification and biomass (Hansen and Nielsen, 2008).

According to Haid and Koss (2001), conventional steam reforming is economically

applied to medium sized methanol plants and the maximum single train capacity is

limited to about 2500 mtpd. On the other hand, pure autothermal reforming (ATR) is

cheapest at capacities of 7000 mtpd (Hansen and Nielsen, 2008). However, it is found

that for mid-size capacities in the range of 2500 – 7000 mtpd, a hybrid two-step

(combined) reforming is the best choice as compared to conventional steam reforming

and pure autothermal reforming only (Hansen and Nielsen, 2008; Nielsen, 2008).

According to Cheng and Kung (1994), the methanol production using steam

reforming is a relatively clean and environmentally safe process. As natural gas is burnt

to produce the heat required for the endothermic reforming reaction, CO2 will be

produced in the reformer furnace combustion zone. The flue gas from the convection

side of reformer contains NOx, CO, CO2, volatile organic compounds (VOC) and

particulates.

In the modern methanol processes, the main environmental objective is to reduce

the CO2 emissions. By reducing the CO2 emissions, the impact of methanol production

on global warming can be greatly reduced. Therefore, CO2 could be recovered from the

flue gas by using a pressure swing adsorption (PSA).

45

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Besides that, since the amount of VOC and particulates is not significant in the

reformer using natural gas as fuel, they do not pose a hazard to the environment (Cheng

and Kung, 1994). The amount of NOx generation depends on how the natural gas fuel is

burnt (Cheng and Kung, 1994). By using a clean fuel gas and controlled combustion,

the concentration of NOx in the flue gas will be low. Table 2.7 shows the typical

contaminants from various sources in a methanol plant.

Table 2.7: Contaminants from Various Sources in a Methanol Plant (Reforming Section).

Methanol Plant Effluents Contaminants

Flue gas in steam reformer CO, CO2, NOx, VOC, particulates

Process condensate Total dissolved solids (TDS), total suspended solids (TSS)

Spent catalyst Various metals

Cooling tower blowdown Total dissolved solids (TDS), total suspended solids (TSS)

2.2.8 Selection of Reforming Technology

There are a number of factors which determines the choice of reforming

technology to be used. These include feedstock composition, capital cost consideration,

environmental constraints, cost of utilities such as steam and cooling water and so forth.

Every technology has its own pros and cons. The choice of reforming technologies used

in this design involves a pre-reforming process followed by a two-step (combined)

reforming technology.

Since the feedstock of natural gas consists of heavier hydrocarbons such as

ethane, adiabatic pre-reforming (APR) is essential to convert all the ethane into a

mixture of methane, carbon monoxide, steam and hydrogen assuming sufficient catalyst

activity. Besides that, steam reforming of natural gas will undeniably continue as the

choice of technology to produce syngas due to its most extensive industrial experience.

The steam reformed gas will enter the ATR (secondary reformer) to be further reformed

to produce syngas of stoichiometric ratio close to 2.0 which is vital for a methanol

production plant. Also, ATR is chosen as one of the technologies due to the much lower

CO2 and NOx emissions (30% and 80% reduction respectively) as compared to SMR. In

addition to that, a low methane slip is achieved whereby most of the methane would be

46

SYNTHANOL SDN. BHD.


converted to syngas. Without a proper stoichiometric ratio of syngas, a low yield of

methanol is obtained together with all side reactions.

Partial oxidation is not selected since this process is much more complex as

compared to SMR and ATR. Also, it is less reliable than SMR and incurs a higher

operating cost (Cheng and Kung, 1994). Not only that, partial oxidation is mainly

applied for heavy oil and naphtha which are not included in the feedstock of natural gas.

On the other hand, GHR was not selected because almost all of the heat from the

high temperature product gas will be used to drive the reforming reaction in the steam

reformer. This configuration is advantageous due to the elimination of fired steam

reformers. However, in our plant, many streams required preheating prior to entering

their respective processes. These preheating took place in the convection section of the

top fired steam reformer. Therefore, the selection of a GHR would have been

inappropriate as many other heat exchanges would be required for the preheating of

various streams. This would have incurred a higher expenditure in terms of capital cost

due to the installation of heat exchanges as well as operating cost due to the utilization

of steam and maintenance of these heat exchangers.

47

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2.3 Methanol Synthesis

In methanol synthesis, syngas that is produced from reforming section would be

used to convert into methanol. The main reaction that occurs in the reactor is presume to

be conversion of carbon monoxide and hydrogen to methanol and reversed water gas

shift reaction. It should be noted that in some literatures, the conversion of carbon

dioxide and hydrogen to methanol is also considered.

CO+2H 2⇌CH 3 OH

CO2+H 2⇌CO+H 2 O

Besides that, there are many side products that are produced in the reactor.

However, only production of acetic acid is taken into consideration in this report.

CO+CH3 OH⇌CH3 COOH

The existing technology of reactor for methanol synthesis was examined and

evaluated based on a few criteria such as feedstock quality, reactor design, economics

and other relative advantages and disadvantages between the reactors.

From literature review, it was found that most of the information on methanol

convertor is associated with the technology of the companies such as stated in Lee et al.

(2007). Therefore, after reviewing these technologies, the summary of the findings were

stated in this report and were presented in Figure 2.14 and Table 2.8. Generally,

methanol convertor system can be distinguished into two types which is the fixed bed

system and three-phase system. Fixed bed system corresponds to a reactor that methanol

synthesis reaction takes place in a fixed bed packed with catalyst while for three-phase

system, conversion of syngas (gas) occur with the aid of catalyst (solid) that is fluidize

in an inert liquid phase substance (Sherwin et al., 1975).

48

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Figure 2.14: Types of reactors available commercially (Shaded box corresponds to company that used these reactors).

49

Methanol Synthesis Reactor

Fixed bed system

Adiabatic

Adiabatic Quench

ICI - Axial radial multibed reactor

Adiabatic in series

Kellogg, Brown and Root offers multiple adiabatic reactor with

interstage cooler

Isothermal

Boiling water reactor (BWR)

Linde, Toyo Engineering, Casale

Tube cooled reactor (TC)

Other variants/ combinations

Combined Converter

(TC+BWR)

Lurgi

Three-phase system / Slurry phase reactor

Chem. System

SYNTHANOL SDN. BHD.


2.3.1 Three Phase / Slurry Phase Reactor

Three-phase system reactors or slurry phase reactor has been acknowledged as the

only alternative to fixed bed reactor type (Hansen et al., 2008). Although developed more

than one decade ago, it has only gained academic and industrial interest in the recent 15

years. Three phases system comprises variants of reactors such as the bubble column,

internal loop airlift, external loop airlift and spherical reactor. Most of these reactors are in

a research stage. Chem. System has commercialized its slurry type reactor which based on

bubble column concept as illustrated in Figure 2.15 (Wang et al., 2007).

In this reactor, syngas is fed in from the bottom of the column and is bubbled

through the hydrocarbon oil which contains suspended catalyst. The reactor occurs when

the gas is in contact with the catalyst. Since methanol synthesis is an exothermic reaction

and therefore the heat is absorbed by the hydrocarbon liquid as sensible heat as well as heat

of vaporisation. The temperature of the liquid is controlled by circulating boiler feed water

(BFW) and steam would be generated. Some hydrocarbon oil vaporised together with the

50

Figure 2.15: Illustration of slurry phase type reactor (Wang et al., 2007).

SYNTHANOL SDN. BHD.


product methanol in the heating process and a condenser is needed to separate methanol gas

and liquid hydrocarbon by lowering the temperature (Sherwin et al., 1975).

One of the main advantages of slurry phase reactor is the efficient temperature

control as explained above. Another benefit of using this reactor is it has been tested large

scale and found feasible. Apart from that, the slurry phase reactor could also be

advantageous in terms of its low pressure drop in the reactor. The need of recompression of

gas due to high pressure drop could increase the operating cost and capital cost of a plant.

Since this reactor portrays high efficiency and conversion in methanol synthesis, thus the

use of catalyst is comparably lower than fixed bed reactors (Wang et al., 2007). Lastly,

another advantage as suggested by Hansen et al. (2008) is that this reactor is absent of

diffusion limitations because of its low catalyst diameter.

In the economical aspect of the reactor, it was found to have controversial findings.

Graaf et al. (1996) found that three phase system has higher annual cost as compared to

fixed bed reactor while Nizamof (1989) of Chem. System found that the cost of slurry

reactors is comparable to fixed bed reactor. However, by considering the more recent

literature and understanding that Chem. System developed the system and might possessed

unintended bias toward the system, therefore, the cost of the reactor is presumed to be

relative higher as compared to fixed bed reactor. Another remarkable disadvantage of this

slurry phase reactor is that multiphase flow behaviour analyses is complex (Wang et al.,

2007). The multiphase flow behaviour in the reactor is greatly influenced by high pressure

and temperature. Therefore extensive study on hydrodynamics, mass transfer and liquid-

solid interaction is still needed.

2.3.2 Fixed Bed Reactor

There are mainly two types of fixed bed reactor namely the adiabatic reactor and

isothermal reactor. Other variant of fixed bed reactors do exist but is rarely used or not

widely used. The adiabatic reactor is divided into two main types which is adiabatic quench

reactor and adiabatic reactors arranged in series with inter-stage cooling. On the other hand,

51

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isothermal reactor consists of mainly boiling water reactor (BWR) and other variants of

steam rising reactors. There are many other variants of these reactor types such as tube

cooled reactor.

2.3.3 Adiabatic Quench Reactor

Quench type reactor generally operate by having a fixed bed which contains catalyst

installed in the vessel (Cheng et al., 1994). ICI has developed a low-pressure quench

converter packed bed that contains of catalyst supported by inert material. Cold fresh or

recycled syngas and is quenched in the reactor which enables the control of temperature in

the converter. The gases are introduced into the reactor by spargers known as lozenge

(Spath et al., 2003). This type of reactors is obsolete in recent days and therefore ICI has

developed an improved version of the reactor known as axial radial concept (ARC)

multiple bed quench reactor. This concept comprises up to five multiple fixed bed reactor

arranged in series in the adiabatic reactor. In this design, the cold syngas is quenched at

different intervals between the packed bed catalysts. An illustration of ICI ARC quench

reactor is shown in Figure 2.16.

52

Figure 2.16: ICI quench reactor (GBH Enterprsise, n.d.).

SYNTHANOL SDN. BHD.


It should be noted that the ARC is commonly used because it is inexpensive as

compared to other reactors (Uhde, n.d.). Uhde (n.d.) has also suggested that this technology

could be applied in plants that consist of surplus of steam in the process.

According to Hansen et al. (2008), the reaction trajectory of the reactor is far less

from ideal and it has relatively bad temperature control as compared to others. Besides that,

quench reactor has long known to have poor mixing since the early development of the

reactor by ICI. This is due to the fact that cold syngas quenched into the bed causing

variation in temperature in the bed. The effect passed down through the whole catalyst bed

leading to sever operating difficulties. In addition to that, the poor mixing which is cause by

‘Cold Core’ effect could affect the conversion of the reactants. This effect occurs due to the

fact that high gas flow rate in portions of catalyst with high voidage and low gas flow rate

in portion which has low voidage. Moreover, this irregular temperature distribution in the

catalyst bed encourages catalyst deactivation and formation of by-products (GBH

Enterprsise, n.d.). Another drawback of this type of reactor is that it requires relatively

more amount of catalyst than any other type of reactors. Therefore, the ARC reactor was

introduced to alleviate these problems (Hansen et al., 2008). However, ARC reactor has

exhibited instability which is indicated by varying inlet and outlet temperature following a

sine wave function. Furthermore, this reactor comprises large number of operating

variables in the reactor and making the process difficult to be optimised and controlled.

Another major disadvantage of this type of reactor is that the multiple beds there is increase

in pressure drop through the catalyst bed making higher rate of syngas recompression and

thus resulting in energy consumption penalty. According to Lou et al. (2005), higher power

consumption is needed in quench type reactor as compared to other type of reactors.

2.3.4 Adiabatic Reactors in Series with Inter-stage Cooling

This type of reactor is a simple reactor with catalyst bed packed in a vessel and

reaction occurs in it adiabatically. Each catalyst layer is placed in separate reactor vessel.

Inter-stage coolers are installed in between the reactor vessels. This type of vessel operates

53

SYNTHANOL SDN. BHD.


at similar concept to the ICI quench with the difference that the catalyst layers are installed

in one single and the inter-stage cooling is by quenching cold syngas in ICI technology.

The recycled gas is fed at the first reactor which increases the kinetic driving force and thus

reducing the catalyst usage relative to quench type reactor. The reactor is spherical in shape

and pressure could be reduced which save cost in material construction as shown in Figure

2.17 (Tijm et al., 2001). The methanol conversion reaction is exothermic and therefore

cooling is required to optimize the reaction before entering another adiabatic reactor. This

design has also been adapter by Haldor Topsoe and Krupp Uhde.

According to (Hirotani et al., 1998), this type of reactor uses less catalyst as

compared to quench reactor but one of the disadvantages is that this type of reactor requires

several high pressure reactor as well as many heat exchangers. This contributes to the

increase in capital cost of the plant. Moreover, the reaction pathway of this type of reactor

is far from the maximum.

54

Figure 2.17: Spherical reactor in series (Cheng et al., 1994).

SYNTHANOL SDN. BHD.


2.3.5 Tube Cooled Reactor

In tube cooled reactor, the catalyst is packed on the shell-side of the reactor. The

feed syngas enters the reactor at the bottom of the reactor and distributed through the tubes

and preheated by the heat of reaction developed on the shell side of the reactor. The syngas

then reached the top of the reactor and diverted to the shell-side as shown in Figure 2.18.

The tube cooled reactor was designed by ICI initially and adapted by Lurgi as an

integrated system with one tube cooled reactor and two boiling water reactor. According to

Uhde (n.d.), this reactor type has low catalyst requirement and the capital cost for this

reactor is low. Besides that it requires less equipment item for this reactor as well as

recovering more heat as compared to quench type reactor. GBH Enterprsise (n.d.) pointed

out that this reactor design resulted in apparent cold and hot region within the reactor thus

leading to rapid catalyst deactivation and high level of by-products. This leads to catalyst

deactivated before reaching its design life and hence replacement of catalyst is required.

This problem not only causes a significant increase in production cost, but also forces the

plant to shutdown abruptly.

55

Figure 2.18: Tube cooled reactor design (GBH Enterprsise, n.d.).

SYNTHANOL SDN. BHD.


2.3.6 Isothermal Boiling Water Reactor (BWR)

The boiling water reactor (BWR) is one of the most commonly available steam-

rising reactors. The other reactors of this kind of reactor would be Toyo MRF-Z reactor and

Linde Steam Rising Converter with internal spirally wounded tubes. However, these

technologies are not commonly used and the disadvantages could not be found as there are

no literatures regarding these technologies. The BWR has a design similar to shell and tube

heat exchanger. The catalysts are packed in tubes and the tubes are immersed in boiling

water. The exothermic reaction in the tube side provides heat to boiling water in the shell-

side. The boiling water absorbs heat and produce steam in the steam drum. The illustration

of BWR is shown in Figure 2.19.

This contributes to good temperature control in this isothermal reactor. The reactor

temperature can be controlled by varying the steam pressure and stable temperature could

be achieved as opposed to quench and tube cooled reactor (DPT, n.d.). This type of reactor

has the most efficient temperature control system as oppose to other reactors (Uhde, n.d.).

These types of reactors are easily controlled as compared to quench type and the reaction of

56

Figure 2.19: Boiling Water Reactor (BWR) design (Rahimpour et al., 2008).

SYNTHANOL SDN. BHD.


rate is close to the optimum reaction rate. These factors contribute to high yield and high

selectivity. According to Lou et al. (2005), the power consumption of the BWR is lower

than other quench reactor. Moreover, this type of reactor has lower operating cost as

compared to other type of reactors. This advantage offset the high capital cost due to the

fact that operating cost runs for the overall plant life while capital cost is only a paid-once

sum. Furthermore, while using a BWR, the catalyst lifespan is longer as compared to a

quench type reactor. Another notable advantage of BWR is the production of steam in the

reactor. The steam generated could be used in reforming section or generate electricity with

a turbine. Last but not least, BWR has lower pressure drop across the catalyst bed as

compared to other reactors. The low pressure drop could minimize the operating cost in

recompression of syngas recycle back to the reactor (Bartholomew, 2006).

However, the disadvantage of this design is the complicated design which

contributes to the high capital cost. This type of reactor has maximum size constrain of 6 m

(Diameter) which corresponds to a single line capacity of up to 1800 t/day.

2.3.7 Reactor Selection

In selection of the suitable reactor type for methanol synthesis process, there are

three main criteria, i.e. temperature control in the reactor, pressure drop in the reactor and

economics, which is needed to be considered (Lange, 2001).

Firstly the temperature control of the reactor must be efficient. This is due to the

fact that methanol conversion is an exothermic reaction and inefficient of temperature

control could lead to temperature rise beyond the design temperature. Excessive heating

could cause severe effect in yield as well as selectivity. For example, excessive heating

cause thermal degradation in catalyst which then lead to low conversion high production of

by-products and reduce catalyst life span in which then, these lead to high production cost

of methanol. At highly elevated temperature, methanation would occur and lead to

catastrophic effect in the reactor since methanation is self-propagate and high exothermic.

Therefore, effective temperature control could prevent methanation in the reactor. As a

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comparison between the commonly used reactors, Boiling Water Reactor (BWR) and

Slurry Phase Reactor has stable and good temperature control whereas quench type and

tube cooled reactor does not.

Secondly, the pressure drop in the reactor does contribute to one of the reasons of

reactor selection. High pressure drop indicates higher rate of compression needed and

hence increase operating cost. Furthermore, high pressure in reactor affects the reaction rate

as well. For instance, BWR and slurry phase reactor both have low pressure drop as

compared to adiabatic quench type reactors.

Thirdly, the technology needs to be evaluated from the economics point of view.

For capital cost, adiabatic quench reactor and tube cooled reactor has notable advantage as

compared to BWR and slurry phase reactor. However, the operating cost of BWR and

slurry phase reactor is much lower as compared to quench reactor.

As a summary of the reactor technology selection, the Boiling Water Reactor

(BWR) was selected as the synthesis reactor in the process due to the fact that it has good

and stable temperature control and leads to high productivity and low by-products

formation. Low by-product formation gives potential advantage over the over reactors as

the minimum treatment is needed before discharging to the environment. Moreover, this

reactor produces steam which could either be superheated to be used in generating

electricity or could be supplied to the steam reforming section. In term of economics, the

operating cost is low and thus relieves the burden over the operating life of the plant.

However, the production capacity of BWR is low (up to 1800 t/day) and therefore two

BWR reactors were used in the design which corresponds to a maximum capacity of 3600

t/day.

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Table 2.8: Summary of Methanol Synthesis Technologies.

Reactor Main Features Advantages DisadvantagesSlurry Phase Reactor(Hansen et al., 2008)(Wang et al., 2007)(Sherwin et al., 1975)(Graaf et al., 1996)(Nizamof, 1989)

Catalyst suspended in hydrocarbon oil (Fluidised bed)

Good temperature control Low pressure drop Low operating cost Relatively less catalyst used

High capital cost Complex multiphase flow

behaviour analyses

Adiabatic Quench(Cheng et al., 1994)(Spath et al., 2003)(Uhde, n.d.)(Hansen et al., 2008)(GBH Enterprsise, n.d.)(Lou et al., 2005)

Up to 5 adiabatic catalyst bed installed in series in a pressure vessel

Relatively cheap Non-ideal reaction trajectory Poor mixing Poor temperature control Formation of by-products Large amount of catalyst

needed Difficult process control and

optimised High pressure drop

Adiabatic series reactor(Tijm et al., 2001)(Hirotani et al., 1998)

Adiabatic packed bed reactor with inter-stage cooling

Less catalyst Large number of HP reactor heat exchangers and pipe cost

Reaction path away from maximum

Tube cooled reactor(Uhde, n.d.)(GBH Enterprsise, n.d.)

Catalyst is packed on the shell-side of the reactor and the reaction preheating the entering syngas feed

Low cost Low catalyst requirement

Rapid catalyst deactivation High level of by-products

Isothermal boiling water reactor(DPT, n.d.)(Uhde, n.d.)(Lou et al., 2005)(Bartholomew, 2006)

The catalysts are packed in tubes and the tubes are immersed in boiling water. The boiling water absorbs heat and produce steam in the steam drum.

Most efficient temperature control

Reaction rate close to optimum High yield and high selectivity Low power consumption and

operating cost Long catalyst lifespan Produces steam on shell-side

Design complication High cost Low capacity

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2.3.8 Technology Evaluation of Catalyst

Catalyst selection is a very important procedure in a chemical plant design since

it has a significant contribution in expenditure. It has a major impact on the plant rate,

plant efficiency and on what is the desired turnaround schedule. Any unforeseen

possibilities of catalyst failure might cause big losses of capital due to inefficient

production.

In selecting a catalyst, several aspects need to be taken into consideration. The

first criteria would include activity, pressure drop and strength. A high conversion is

usually required which is as closest to the equilibrium. A low pressure drop ensures

higher efficiency and plant rate. High strength catalysts are desirable since zero damage

is wanted during the loading process and a high rate of reaction needs to be sustained

through a stable pressure drop with time.

More factors that need to be considered for catalyst selection are activity

retention, selectivity, poison resistance and heat transfer. The longer lifetime the catalyst

has the better and more cost efficient it is. Catalysts are more efficient when they are

selective since they will only catalyse the required reactions and not produce other by-

products (Hawkins, 2011). Heat transfer within the catalyst is important since the rate at

which gas molecules diffuse onto the catalyst surface for adsorption will affect the

overall rate of reaction.

The structure of the catalyst is also governs the catalyst efficiency. The fluid

flow through the catalyst bed depends on the shape and size of the catalyst and the

mechanical strength ensures the lifetime is long enough. A high surface area and rightly

selected chemical components will ensure optimal activity and selectivity. The support

of the catalyst should possess high enough surface area for the active components to be

evenly distributed to avoid undesired sintering (Richardson, 1989).

In methanol synthesis, catalyst selection is crucial since hydrogenation of carbon

monoxide and carbon dioxide favors higher alcohols over methanol as products and

dimethylether may also form. Currently, catalysts that allow the production of nearly

pure methanol from synthesis gas at the low pressure of less than 100 atm are available.

These contain copper and a mixture of oxides for instance, ZnO/Al2O3 or ZnO/Cr2O3

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and most industrial methanol synthesis has largely been carried out on these two types

of catalyst (Hansen, 2008). As pointed out by Andrew et al. (1980), the main features of

methanol-synthesis catalysts are:

(a) Fairly good hydrogen activation ability, which is usually not considered to be a

limiting factor in the reaction.

(b) Activation of CO without dissociation (cleavage of a C-O δ-bond: 360 kJ/mol), as

otherwise methanation occurs.

(c) Absence of undesirable support components, such as active alumina (excessive

dehydrating activity), nickel and iron impurities (excessive hydrogenation activity), and

sodium impurities (excessive alkalinity).

Table 2.9 shows a review of some of the proposed catalysts for methanol synthesis.

Catalyst Advantages Disadvantages

ZnO/Cr2O3Highly resistant to catalyst poisoning, especially towards sulphur.

Requires high temperature and pressure, currently obsolete, not in use industrially, no longer economical

Cu/ZrO2

Methanol synthesis reaction rate increased, higher adsorption capacity of carbon oxides

Slow reverse water shift gas reaction, much less CO produced.

Cu/ZnO

Low temperature and pressure, Reduction of compression and heat exchange duty in recycle loop. improved selectivity by suppressing production of light hydrocarbons

Deactivates quickly as temperature increases

Cu/ZnO/Al2O3

High activity, very good selectivity,long-term stability, and favorable production costs, most cost effective catalysts, easily available on the market, most exclusively used methanol synthesis catalyst,high poison durability relatively low reaction temperature and pressure

Activity loss with water, sintering at high temperature.

Cu/ZnO supported on Pd

High activity, long lifetime, high selectivity

Not readily available commercially

Pt-based catalyst Very active and selectiveUse of noble metals not commercially feasible

On the basis of the above comparison in Table 2.9, the catalyst selected for

methanol synthesis will be the Cu/ZnO/Al2O3 system. Table 2.10 below shows the

different productivities of methanol using different compositions for the Cu/ZnO/Al2O3

system at different operating conditions.

61

Table 2.9: Review of different catalyst for methanol synthesis (Mäyrä et al.,2008).

SYNTHANOL SDN. BHD.


Table 2.10: Productivities of methanol using different compositions for the catalyst (Herman et al., 1979).

In typical industrial operating conditions (70 – 100 bar, 220 – 280 ˚C, 30.000 –

40.000 h-1 flow rate) raw methanol (80% MeOH, 20% H2O) is produced in modern

plants using Cu/ZnO/A12O3 catalysts. Major impurities are higher alcohols, methyl

formate and hydrocarbons. The production of higher alcohols is greatly suppressed by

CO2 in the feed, as no chain-growth mechanism operates.

Catalyst life is directly proportional to the ability of the catalyst to absorb

poisons in the feed. The zinc oxide component is the best absorbent, as shown by a

thermodynamic analysis of the relative ease of formation of chlorides and sulfides.

Poisoned catalysts show ZnS formation. In order to guarantee good sulfur absorption it

is therefore necessary to have a catalyst formulation containing a high surface area of

exposed free zinc oxide (this is more desirable for water-gas shift (WGS) catalysts).

Halogen induced sintering (through formation of volatile copper chloride) is retained

being one of the chief causes of copper crystal growth in methanol and shift catalysts.

(Bart et al., 1987)

A good methanol catalyst formulation may therefore be composed of an

adequate surface area (typically 50Å particles) of copper and zinc oxide (for

chemisorption and catalysis) and a finely dispersed (20 Å) refractory support (e.g. Al2O3

or ZnAl2O4) to counteract thermally induced sintering. High methanol selectivities are

best achieved using ZnA12O4 instead of Al2O3 but the most available catalyst in the

market is the Al2O3. Current drawbacks of industrial Cu/ZnO/A12O3 catalysts, however,

comprise a relatively important drop in activity in a 3-year production run (75%) and

varying catalyst quality.

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2.4 Product Purification

2.4.1 Single and Two-column Distillation Column

In early stage of distillation of crude methanol, single distillation column

operating in low pressure was used to achieve the objective. However, with the

significant rise in energy cost in mid-1970s, the distillation column was kept modifying

up to different designs to achieve higher energy efficiency with least energy consumed

(Douglas, 2006). The distillation process was also optimized to improve the process in

more economical and sustainable approaches. The schematic diagram of a single

distillation column is shown in Figure 2.20.

Figure 2.20: Schematic diagram of distillation column (Scott, 1977).

Currently, the most conventional method of distillation used in industry will be

the two-column methanol distillation scheme which basically comprises topping and

refining columns. The typical arrangement and schematic diagram of two-column

distillation column is shown in Figure 2.21. Both of the distillation columns are

operated at approximately atmospheric pressure (~1 bar). Eventually, 98.5% of

methanol from methanol synthesis process can be recovered through the two-column

distillation scheme (Uhde, 2011).

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Theoretically, the function of topping column is to remove light-end components

which have lower boiling point than methanol such as dissolved gases (CO, CO2, H2,

etc.), dimethyl ether, methyl formate, and acetone. Likewise, the function of refining

column is to remove the heavy-end components with higher boiling point than methanol

such as include water, higher alcohols, long-chain hydrocarbons, higher ketones, and

esters of lower alcohols with formic, acetic, and propionic acids. All the heavy and

light-end components are removed from the distillation columns as wastewater and tail

gas respectively in the methanol production process. The purified methanol obtained

eventually will be sent for storage and utilized in other industries. The essentially pure

wastewater will be discarded or reused within the process whereas the tail gas with

certain amount of different gases will be further separated as fuel for reformer or other

heating equipment (Siemens, 2007).

Basically, the crude methanol feed from methanol synthesis process with the

temperature and pressure of 40°C and 5 bar is fed into the ¼ (34th trays) from bottom of

topping column consisted 42 trays in total. The light-end products with temperature of

70°C will be distillated in condenser on top of column to 45°C and thus to be burned off

by mixing with reformer fuel. Some of the bottom products which are predominant in

liquid methanol leaves at 80°C and 1.65 bar will be reboiled up to 88°C and the left

liquid products consists of predominant methanol will be pumped into refining column

with the pressure of 3.11 bar for further distillation (Hawkins, n.d.; Pinto, 1980).

In the refining column, the liquid products are further distillated at 81 °C and a

methanol product with 99.99% minimum purity and low impurities can be obtained

which satisfies the specification required (Hawkins, n.d.). The methanol products are

condensed and routed to storage tank at normal conditions of 20°C and atmospheric

pressure which are defined by World Health Organisation (WHO) (Organisation, 2011;

Trifiro, 2009). The bottom product in 125°C and 2.30 bar which is predominant in water

will be reboiled to 130°C and the remaining bottom products with the methanol content

of 0.1% will be used as water source within the process by cooling down to desired

temperature (Hawkins, n.d.).

Due to the presence of water and ester in crude methanol stream, corrosion of

equipment might occur during distillation and storage stages. Besides the use of

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corrosive-resistant materials, dosing of aqueous caustic soda (NaOH) in crude methanol

before flowing into distillation column is preferable by making the acidic feed stream

slightly alkaline (pH > 7) to prevent corrosion of distillation column as well as the

piping during the operation (Fiedler, 2005). The amount of NaOH required is basically

one litre of 2% per tonne of methanol (GBHE Entreprise Ltd., n.d.).

Figure 2.21: Schematic diagram of two-column distillation column (Cialkowski, 1994).

2.4.2 Optimization of Process Technology

Due to the consideration of energy efficiency, the potential of mass and energy

savings provides a significant aid to achieve the objective. One of the methods is to

introduce a series of multi-effect distillation columns with efficient heat integration

between columns which can have significant lower mass and energy requirements as

compared to conventional two-column distillation scheme. For instance, a five-column

scheme with addition of medium-pressure column after original higher-pressure column

can significantly reduce the load of higher-pressure and atmospheric columns by 30%.

Besides, the economic analysis on energy consumption of five-column scheme

shows a reduction of 33.6% as compared to four-column scheme (Zhang, 2010). The

more the distillation columns being introduced, the higher methanol recovery and lower

steam consumption can be obtained. For example, 99.5% of methanol recovery and

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SYNTHANOL SDN. BHD.


reduction of steam consumption by 0.75 tonnes per tonne methanol as compared to two-

column scheme can be obtained by adopting four-column distillation scheme (Uhde,

2011). The following table (Table 2.11) shows the comparison of condenser and

reboiler duty as well as the steam consumption between four and five-column schemes.

Table 2.11: Comparison of calculation results between different schemes (Zhang, 2010).

As shown in Table 2.11, the total heat requirement as well as the consumption of

cooling water and steam of five-column scheme show a lower value as compared to

four-column scheme with the approximately same purity of methanol obtained

eventually.

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Table 2.12: Specification of Grade AA methanol (Pound, 1998).

Besides of the consideration of energy efficiency, the quality of product

dominates the design of the distillation section. In order to produce high purity

methanol which meets the US federal specification O-M-232K Grade “AA” with 99.85

wt% purity (Table 2.12), optimisation of process design is done and finally a three-

column distillation scheme with addition of recovery column is introduced in two-

column distillation scheme to achieve this objective. The purity of methanol obtained

from this scheme can be achieved up to 99.99% (Zhang, 2010). The schematic diagram

of three-column distillation is shown in Figure 2.22.

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Figure 2.22: Schematic diagram of three-column distillation (Douglas, 2006).

Due to the ethanol build up in the middle of refining column because of the non-

ideal behaviour of ethanol in presence of water. Ethanol is more volatile than methanol

at higher water concentration in stripping section of refining column. When the stream

moves upwards results in decrease in water content and methanol dominates the higher

volatility. As a result, the ethanol reaches maximum concentration in the middle of

refining column (Uhde, 2011). Thus, the recovery column plays the role of withdrawing

the middle boiling impurities (principally ethanol, but also higher alcohols, ketones and

esters) as side stream, which is called as fusel oil, that is basically used for primary

active ingredient in all alcoholic beverages (Hori, 2003; Zhang, 2010). It can be used as

chemicals for flavour and fragrance manufacturing. Apart from commodity industry,

fusel oil can be used for phosphoric acid purification by wet method in chemical

manufacturing industry (Kucuk, 1997). For certain recycling of wastewater, a

significant amount of acetic acid will be obtained and thus can be further extracted out

from water for usage in chemical industry as derivatives. The largest consumption of

acetic acid will be the manufacture of vinyl acetate monomer (VAM) which can be used

in production of emulsions such as base resin for water-based paints, adhesives, paper

coatings and textile finishes (ICIS, 2011).

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Figure 2.23: Configuration of two-stage separator and distillation column (William, 2010).

Another alternative technology of distillation section will be a combination of

separators and distillation column which is shown in Figure 2.23. Due to the simplicity

of process, topping column suggested in two-column scheme can be substituted with

separators operating at different pressures and thus the cost saving can be achieved as

well by substitution of cheaper equipment due to simple construction and smaller

dimensions.

Since the high pressure of crude methanol obtained from methanol converter, a

high pressure separator is required for primary separation of light-end gases from liquid

products. The stack gas with trace amount of moisture will be recycled back to

methanol synthesis process due to the significant amount of gases which can be reused

within the process to increase product yield and improve the process sustainability.

Because of the requirement of high stream pressure in order to recycle into converter,

the high pressure separator is chosen instead of reducing the pressure and being

separated in low pressure separator. Instead of full recycle of gases, some of the gases

will be purged off from the process as the waste and mixed with reformer fuel to be

burned off. The reasons will be to sufficiently control the flow rate of recycled gas and

remove the inert substance such as N2 from product stream to avoid accumulation which

affects the performance of methanol converter.

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Due to high operating pressure, some of gases might retain in the liquid stream

and thus another separator has to be installed for further separation. Since the trace

amount of useful gases which are considered not economical-friendly by recycling back

into process, low pressure flash tank is introduced as secondary separation in terms of

safety and economical consideration. The liquid product from high pressure separator

will be expanded through a throttling valve to the pressure which is consistent with the

operating pressure in flash tank. Flashing of methanol or other substances from the

expansion will be occurred and some of moisture will be separated as well with the gas.

Thus, installation of demister pad is essential to retain the 99% moisture from gas and

the all the residual gases will be flowed through as overhead product and mixed with

purged gas obtained in high pressure separator. Before feeding into distillation column,

the pressure of pure liquid product will be reduced using pressure regulator in order to

fit the operating condition for efficient distillation.

The high and low pressure separators are crucial in the process as the adverse

effect of blanketing of inert components in condenser due to significant amount of light-

end gases fed into distillation column can be eliminated and thus the distillation column

can be operated sufficiently (William, 2010). In distillation column, the methanol

product will be separated as overhead product with the minimum purity of 99.85%

whereas most of the water and acetic acid will be separated as bottom product which

will be reused as feed water within the process. Extraction of accumulated acetic acid is

required after certain period of time for other purposes. The operating conditions of

reboiler and condenser depend on the design and the methanol will go through a series

of condensation and vaporization within the distillation column. Eventually, the

methanol product will be routed to storage tank with the operating condition at 1 bar

and 20°C.

Plate Contactors

The main requirement of a tray is that it should provide intimate mixing between

the liquid and vapor steams and suitable for handling desired rates of vapor and liquid

without excessive entrainment and flooding. The arrangements for the liquid flow over

the tray depend largely on the ratio of liquid to vapor flow. There are three types of

liquid flow configuration namely cross-flow, reverse and double-pass as illustrated in

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Figure 2.24 respectively. Reverse flow is more suitable low liquid-vapor ratios, whereas

double-pass configuration is used to handle high liquid-vapor ratios (Baer, 2011).

The most common type of plate contactors used for tray distillation column is

cross flow plate, which consists of the bubbling area and vertical channel ‘down-

comers’ providing good length of liquid path, hence enhance mass transfer (Sinnot et

al., 2009). Liquid descending from plate to plate via ‘down-comers’ enters bubbling

area, a pool of liquid is retained on the plate by an outlet weir. There are three principle

types of cross-flow plate used in industry which is sieve plates, bubble-cap plates and

valve plates. Valve plates can be further differentiated into two categories namely

floating-cap plates and fixed valve plates (Sinnot et al., 2009).

Figure 2.24: Arrangement for liquid flow over a tray (Coulson et al., 1991).

Sieve Plate

Sieve plates are also known as perforated plates is the most commonly used and

simplest type of cross-flow plate. The liquid flows across the tray and down the

segmental down-comer where vapor passes up through perforations in the plate. The

velocity of the up flowing gas keeps the liquid from descending through the

perforations. However, the liquid will somehow weep through the perforations at low

gas velocities due to absence of positive vapor-liquid seal in the plate. Thus, the plate

efficiency will be affected by the weeping effect (Coulson et al., 1991). The operation

of sieve plates is shown in Figure 2.25.

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Figure 2.25: Operation of Sieve Plates (Norrie, 2010).

Bubble-Cap Plates

The bubble cap distillation plates are flat perforated plates with risers (chimney-

like pipes) around the holes, and caps in the form of inverted cups over the risers. The

main advantage of this plate design is that a liquid level is maintained on the top of the

tray at all vapor flow rates as the vapor from underneath the tray pushed through the

bubble cap. Therefore, bubble-caps have good turn down performance at low flow rates

(Baer, 2011). Nevertheless, this is the most costly and complex tray design. The

operation of bubble-cap plates is illustrated in Figure 2.26.

Figure 2.26: Operation of Bubble-Cap Plates (Norrie, 2010).

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Valve Plates

Valve trays may be regarded as a cross between bubble-cap and sieve plates

which possess similar design characteristics of both. Floating-cap valve plates are

essentially sieve plates with large diameter holes covered by movables which lift as the

vapor flow increases (Baer, 2011). For fixed valve plates, it is somehow similar to a

sieve plate but the holes are only partially punched out such that the hole remains

partially covered. Typical operation of the valve plates is shown in Figure 2.27.

Figure 2.27: Operation of Sieve Plates (Norrie, 2010).

Summary of Plate Types

Table 2.13: Comparison of Plate Type (Maloney, 2008).

Sieve PlatesBubble Cap

TraysFixed Valve

PlatesFloating-Cap

ValveCapacity High High High High to very highEfficiency High High High High

Turndown

About 2:1. Not generally

suitable for operation under variable loads

About 5:1.

About 2.5:1. Not generally

suitable for operation under variable loads

About 4:1 to 5:1. Some special

designs achieve 8:1 or more

Entrainment Moderate Moderate Moderate ModeratePressure drop Low Highest Moderate Slightly higher

Cost Low Highest Low About 20% higherMaintenance Low Moderate Low Moderate

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Fouling Tendency

Low to very low Low Low to very low Moderate

Effects of Corrosion

Low Low Very low Moderate

Main applications

Most columns when turndown is not critical

High fouling and corrosion potential

Most columns

Most columns when turndown is not critical

High fouling and corrosion potential

Most columns Services

where turndown is important

Referring to Table 2.13, cost, capacity, operating range, efficiency and pressure

drop are the dominant factors to be considered for the selection of suitable plate type for

a distillation column. Bubble-caps plates are rarely used for new installations on

account of their high cost and pressure drop. In addition, bubble-caps will contribute to

large hydraulic gradients across the column (Coulson et al., 1991). Bubble-caps are only

capable to handle very low liquid rates with low reflux ratios. Due to limitations

mentioned earlier and high cost requirement, bubble-cap plate is the least preference

technology as compared to others. Valve tray offers advantages over bubble-cap and

sieve plates in terms of economical and operational as shown in Table 2.13. However,

due to the proprietary nature of this plate type, information on the design and

performance can only be estimated from published literature. The valve plates are

usually designed by the manufacturer (Coulson et al., 1991).

Sieve plates are deemed to be the most suitable plate type for methanol purifying

distillation column. Sieve trays offer several advantages over bubble-caps and valve

plates such as the simplicity of technology and low installation and operating cost

requirement. The pressure drop for sieve plate is lower as compared to bubble-cap and

valve plates. On top of that, the fundamentals are well-established and hence entailing a

lower risk in the distillation operation (Coulson et al., 1991). Most importantly, sieve

plates experience low corrosion effect and have low tendency to fouling seeing as the

components in the distillation column contain methanol and acetic acid which are

corrosive substances.

Packed Column

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Packed column (shown in Figure 2.28) is an alternative technology for

distillation in which the cylindrical shell of the column is filled with some form of

packing providing large interfacial area for diffusion. The packing may consist of rings,

saddles, or other shaped particles. In packed columns, vapor flows steadily up and the

reflux steadily down giving a true countercurrent system in contrast with tray

distillation column where the process of enrichment is stage wise. Moreover, the

performance of a packed column is dependent on the gas-liquid distribution throughout

the packed bed (Coulson et al., 1991).

Figure 2.28: Packed distillation column (Norrie, 2010)

2.4.3 Selection of Process Technology

In this design project, the selection of a high pressure separator followed by a

low pressure flash tank with the principle of gravitational settling due to density

difference was done. This is followed by a distillation column (William, 2010). The

topping column is replaced by a flash tank due to the redundancy in operation of the

distillation column. In our situation, there are no volatile matters such as aldehyde or

ketone present in the stream and thus the distillation does not apply for this stream

separation. In general, acetone is taken as a key design component in order to design the

topping column with respect to the reflux ratio, number of stages etc. However, the

absence of this substance makes the selection of distillation column inappropriate

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(Cialkowski, 1994). Moreover, the mass and energy balances around the condenser and

reboiler could not be performed as well.

Due to these constraints mentioned, the selection of two-column distillation

scheme was eliminated. The installation of topping column unit becomes redundant

because its real function is not being utilized in our application and this enables us to

decrease the capital investment involved.

The multi-column scheme was eliminated despite its advantage of 30% energy

saving by adopting four-column scheme as compared to two-column scheme (Zhang,

2010). This was due to the fact that this configuration poses a potential of higher capital

cost since more distillation columns are being installed as shown in Figure 2.29. The

energy consumption of second column (C2’) accounts for around 40% of total

consumption for distillation process and it demands great steam consumption in its

reboiler. In addition, the rigorous requirement of methanol content in waste water B5’

has made the current scheme not sustainable due to the difficulties in methanol

separation in atmospheric column C4’ and recovery column C5’ (Zhang, 2010).

Figure 2.29: Four-column distillation scheme (Zhang, 2010).

Plate distillation column is more preferable as compared to packed column for

methanol purifying due to several limitations of packed column. Plate columns can

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handle a wider range of liquid flow rates without flooding as compared to packed

column, seeing that packed columns are not suitable for very low liquid rates. Besides

that, man-holes are provided in plate columns for the ease of maintenance. In packed

columns, packing must be removed before cleaning. Moreover, the design procedure for

plate distillation columns is more well-established with greater assurance as compared

to packed columns. The uncertainty in maintaining a good liquid distribution throughout

a packed column under all operating conditions is the main drawback of this technology

(Sinnot et al., 2009).

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3.0 Process Synthesis and Process Flowsheet Development

3.1 Development of Flowsheet Structure

The process synthesis and flowsheet development was carried out by taking a

basis of 106 metric tons per year of methanol to be produced. Natural gas was fed at

about 635200 metric tons per year into the process line. The major steps through which

the natural gas goes through includes desulphurization, pre-reforming, steam-reforming

and autothermal reforming. This results in the production of synthesis gas which is then

sent to the methanol conversion section. The converter exit stream enters a flash

separator where unreacted and other inert gases are separated from the liquid product

which consists of methanol, acetic acid and water. A refining column is then used to

produce methanol of the required purity, 99.85%. The methanol product is eventually

sent to a storage tank.

The presence of H2S and any other sulfur compounds are undesirable in the

feedstock as these compounds can cause corrosion and hydrogen embrittlement in

certain metals which will reduce the heating value, thus affects the quality of the natural

gas (Hairmour et al., 2005). Catalyst poisoning is also a major possible consequence. It

is therefore of high significance to carry out desulphurization of natural gas. This is

done by feeding the preheated and compressed natural gas into a desulphuriser unit

packed with the sorbent, Sulfatreat.

It is important that the steam to carbon ratio of the syngas entering the reactor is

high enough to approach thermodynamic equilibrium and to reduce side reactions. A

saturator is therefore used to increase the water content of the desulphurised gas before

being sent to the reforming section.

After this, the process stream will then enter a pre-reformer. A pre-reformer is

installed to ensure complete conversion of all higher hydrocarbons present, namely

ethane, in the natural gas feed. Pre-reforming is a necessary step to ensure the

prevention of carbon formation known as hot banding and hot spots in the subsequent

unit operations. It also ensures total removal of any traces of sulphur in the natural gas

feed (Christensen, 1996). Heating utility is required to preheat the saturated natural gas

feed before entering the pre-reformer, which operates adiabatically. The effluent of

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APR will then split into SMR and ATR with a percentage flowrate of 45% and 55%

respectively.

The effluent from the APR will then enter the combined reforming process. The

main advantage of this reforming configuration is the bypass of feed gas across the

steam reformer. Here, only 45% of this effluent will be routed through the SMR

whereas the remaining 55% will be bypassed to the ATR. The bypass stream is

necessary in order to ensure that the overall process steam consumption in the SMR is

roughly halved. Also, the size and the heat load of the SMR would be reduced as the

total flowrate of feed gas into the SMR has been greatly decreased. Evidently, the SMR

load was found to be reduced to 70% of that required by a conventional SMR, hence

decreasing the amount of fuel needed (Uhde, 2006). This would subsequently lead to a

reduced energy requirement and hence a lower investment of the SMR.

The combined reforming process is also beneficial to produce the desired quality

of synthesis gas with a stoichiometry ratio of close to 2.0. This is achieved by attaining

a simple combination of the H2-rich syngas from the SMR with the CO-rich syngas

from the ATR. Other than that, this ratio can also be optimized by adjusting the oxygen

to carbon ratio into the ATR. This stoichiometry ratio is crucial in order to achieve the

highest possible yield in the downstream methanol synthesis process.

On the other hand, a high steam to carbon ratio in the combined reforming

arrangement reduces the formation of soot and methane slip in the synthesis gas. A

methane slip as low as 1 mol% can be attained and this will subsequently increase the

carbon efficiency and thus, enhancing the yield of methanol (Uhde, 2006). The

decrement in soot formation could potentially reduce the chances of carbon deposition

on the catalyst (Petersen et al., 2004).

The catalyst used in the reforming section is nickel impregnated on Al2O3

support (Petersen et al., 2004). Methane is known to be a very thermodynamically

stable molecule even at high temperatures. Therefore, the catalyst is needed to reduce

the operating temperature and hence decrease the tube stresses resulting from high

pressures and temperatures (Logdberg and Jakobsen, 2010).

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Syngas from reforming section needs to be converted to methanol. The

production capacity is affected by the type of reactor used. Therefore, two Boiling

Water Reactors (BWRs) are used to boost the capacity due to the fact that each BWR

could only produce 1800 t/day.

Besides that, the reactors need to be maintained at quasi-isothermal conditions

for a high yield and low by-product formation. Therefore, boiling water is fed to the

shell side of the BWR to cool the reactor and produce saturated water and steam. A

steam drum is needed to separate the mixture of saturated water and steam in order to

produce pure saturated steam.

Furthermore, there are unreacted reactants and products in the outlet of the

reactor due to the fact that the reactions in the reactor are reversible. Consequently, the

reactants and products are separated in a high pressure separator in order to recycle the

gaseous reactants back to the main process stream and the bottom liquids are sent to

purification.

The crude methanol stream contains acetic acid, methanol and dissolved gasses

at high pressure. The crude methanol needs to be further separated from the gases at low

pressure and therefore, a letdown valve is used to letdown the pressure and then it enters

the letdown vessel to separate the remaining gases in the stream before it could be

further purified in the distillation column.

In order to achieve methanol product purity of 99.85%, a distillation column is

essential to separate the light and heavy components from the letdown vessel effluent.

Methanol exits as top product from the column whereas the heavy components (acetic

acid and water) leave the refining column as wastewater which will be appropriately

treated in a waste water treatment plant before being released to the environment. A

reboiler is placed at the bottom of the refining column to provide heat for vaporization

to generate vapors which will be channeled back into the column to drive the distillation

separation.

For the operating column, a total condenser is used so that methanol vapor can

be fully recovered and is then sent to a reflux drum. A portion of the condensed

methanol is refluxed back into the column to enhance the separation efficiency. Due to

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the high volatility of methanol liquid at high temperature, an additional condenser is

placed after the reflux drum in order to further condense the liquid methanol product to

a lower temperature which is feasible for storage. The liquid methanol product with

purity of 99.85% is then sent to a storage tank.

3.2 Reaction

In the desulphurisation section, H2S in the feedstock natural gas reacts with the

iron oxide mixture in the Sulfatreat sorbent producing sulphur, water and iron sulphide.

The reaction equations are shown in equations 1 and 2 (Svärd, 2004). The operating

pressure in the reactor is 40 bar and it operates adiabatically.

Fe2 O3+ H 2 S →2 FeS+S+3 H 2 O(Equation1)

Fe3 O4+H 2 S → 3 FeS+S+4 H 2O( Equation2)

In the pre-reforming section, the higher hydrocarbon, which is ethane in this

case, will be hydrocracked into carbon monoxide and hydrogen gases (Equation 3).

Besides that, methanation of carbon monoxide (Equation 4) and water gas shift reaction

(Equation 5) will also occur simultaneously in the APR. All three reactions as shown

below will be carried out at 500℃ and 36 bar. The hydrocracking of ethane is an

endothermic reaction whereas both the methanation of carbon monoxide and water gas

shift reaction are exothermic. In overall, the whole reaction process in the APR is

slightly endothermic and will cause a drop in temperature of the process stream.

C2 H 6+2 H 2O →2CO+5 H 2(Equation3)

CO+3H 2⇌CH 4+ H 2O(Equation 4)

CO+ H 2O⇌CO2+H 2(Equation5)

In the SMR, the feed undergoes both steam reforming as well as water gas shift

reactions simultaneously. The reactions occurring in the reactor tubes were assumed to

be non-isothermal processes. The process occurs at an operating temperature of 850 ºC

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and pressure of 31 bar. The heat required in the steam reformer is provided by the

combustion that occurs in the firebox.

CH 4+H 2 O⇌CO+3 H 2(Reforming reaction)(Equation 6)

CO+ H 2O⇌CO2+H 2(Water−gass hift reaction)(Equation 7)

The above reactions (Equations 6 and 7) take place whereby methane is partly

converted into carbon monoxide and hydrogen. Some of this carbon monoxide will then

react with water to form carbon dioxide and more hydrogen. The effluent from these

reactions will enter the ATR for further reforming.

There are three main zones in ATR namely the burner, combustion and catalytic

zones. The burner provides a good mixing of the feed gas and oxygen.

CH 4+12

O2⇌CO+2 H 2 ∆ H ro=−35.67 kJ /mol (Equation8)

2 H 2+O2⇌ 2 H2 O ∆ H ro=−483.66 kJ /mol (Equation 9)

The temperature of gas in the combustion chamber is fixed at 1150 ºC (Pina and

Borio, 2006). In the combustion zone, a literature value of 97% conversion of methane

to CO was assumed (Vernon et al., 1990). Here, methane and hydrogen react with

oxygen respectively to produce a combination of carbon monoxide, hydrogen and steam

as shown in equations 8 and 9.

This gas will then be channeled to the catalytic zone whereby steam reforming

and water gas shift reactions take place concurrently. The chemical reactions for both

processes are shown in equations 6 and 7 respectively. Unconverted methane from the

combustion zone will be further reformed to produce synthesis gas which has a

stoichiometry ratio of 1.85. This value is consistent with the reported literature value of

2.0.

In the methanol synthesis section, syngas from the reforming section is reacted

over a Cu/ZnO/Al2O3 catalyst producing methanol, water and by-products. The

reactions occur at 90 bar and 220℃ in the reactor. The reactions are stipulated as below

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(Equations 10 – 12). The overall reaction is an exothermic reaction and thus heat is

produced in the reactor.

CO+2H 2⇌CH 3 OH (Methanol Production)(Equation10)

CO2+H 2⇌CO+H 2 O(Reversed Water Gas Shift ) ( Equation11 )

CO+CH3 OH⇌CH3 COOH (By−product formation)(Equation12)

There are other proposed reactions that occurred in the reactor such as formation

of methanol by carbon dioxide and hydrogen as well as formation of Dimethyl ester

(DME) as by products. However, only Equation 10 – 12 is considered.

3.3 Separation

3.3.1 ATR Effluent

The synthesis gas from the ATR, which is at a high temperature of 1000ºC, will

undergo a heat exchanger and a waste heat boiler placed in series whereby the main

process stream would be cooled to 80ºC and 29 bars. This stream would then enter a

knock-out drum where 90% of water is separated and recycled back to the saturator.

3.3.2 High Pressure Separator

In this section, unwanted materials are separated in two separators operating at

high and low pressures respectively.

The products from methanol reactor are routed into a high pressure separator and

most of the light-end gases are separated at the top whereas the mixtures of liquid gas

products are separated as bottom products. Most of the light-end gases will be recycled

into methanol reactors whereas the remainder will be purged off. The bottom products

are expanded to a lower pressure through a let-down valve and further separated in a

low pressure separator (letdown vessel) in which the residual gases are separated as

stack gas whereas the liquids are separated as bottom product again.

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3.3.3 Letdown vessel

The pressure of purged gas from high pressure separator will be regulated to fit

the stack gas pressure from the letdown vessel. Both the gases will be mixed and routed

to the reformer furnace to be burnt off. The bottom liquids retained from second

separator will be pumped into the distillation column to separate methanol from the

other substances.

3.4 Recycle

3.4.1 Desulphurization Unit

In the desulphurization section, the spent products are both safe and stable

(Braga, 2004). It can be recycled or disposed directly to landfill without any need of

special handling. But the recycling requires additional equipment, handling and extra

cost. Thus, in this methanol plant, the spent product will be disposed of.

Nickel-based catalyst which is used in the packed bed adiabatic pre-reformer

will be spent once the catalyst bed reaches the breakthrough. The spent catalyst will be

disposed of as well since the regeneration technology associates with high capital and

operating cost.

3.4.2 Methanol Synthesis and Methanol Purification

In methanol synthesis section, the deactivated catalyst used in the reactor is sent

to supplier for regeneration of the catalyst. Besides that, the products from the reactor

contains high amount of unreacted reactants due to the reversible reactions in the

reactor. Therefore the reactants are separated using High Pressure Separator from the

crude methanol and recycled back to the reactor. This decreases the feedstock (natural

gas) consumption rate considerably and thus reducing the depletion of fossil feedstock.

A mixture of H2, CO, CH4, CO2, N2, CH3COOH, H2O and CH3OH from the high

pressure separator as well as the let-down vessel would be purged to the firebox in the

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SMR in order to assist the combustion of natural gas fuel and air. By implementing this

method, a significant amount of natural gas that is used as fuel can now be saved and

utilized as feedstock for the synthesis of methanol. Besides, the purged gas obtained can

be said to have no economic value for recycling into the process.

In the distillation column, cooling water used to cool the methanol product will

be recycled back to the cooling tower. Besides that, saturated steam generated from

steam drum in methanol synthesis operation is channeled to the distillation reboiler to

provide heat for vaporization. Then, the cooled water produced from saturated steam in

the reboiler could be recycled back to the cooling tower and reused elsewhere in the

operation.

3.5 Overall Conversion and Yield

3.5.1 Overall conversion

The overall conversion is based on the amount of carbon in the feed and the

product. The methanol product has an atomic carbon flow rate of 4488.62 kmol/hr. All

carbon components in the feed contribute to the formation of methanol and these

include methane, ethane as well as carbon dioxide. The amount of carbon in the feed

was 4882.5 kmol/hr

Thus, the overall conversion is calculated as follows:

Overall conversion=Amount of carbon ( kmol

hr )∈the methanol product

Total carbon( kmolhr )∈the feed

x100

This gives a value of 91.93 %.

3.5.2 Yield

The yield is calculated using the following formula:

Yield= Mole of desired product formedMoles that would have formed if there was no side reactions

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From the process flow diagram, the final stream 7.6 to the storage tank contains

143635.8 kg/hr of methanol which is an equivalent of 4488.62 kmol/hr. A side reaction

of methanol and carbon monoxide led to the formation of acetic acid, decreasing the

amount of methanol being produced. Without this reaction, the amount of methanol that

would have formed would be the same as the amount of acetic acid that has actually

been produced (reaction between methanol and carbon monoxide is equimolar). The

amount of acetic acid formed was calculated to be 10.442 kmol/hr.

Hence, a yield of 99.77 % is obtained for the overall methanol plant.

Mass flowrate of methanol in stream 7.6 = 143635.8 kg/hr

Mr (methanol) = 32

Molar flow = 143635.8

32 = 4488.62 kmol/hr

Mass flow of acetic acid produced = 618.42 + 8.09 = 626.51 kg/hr

Molar flow of acetic acid = 10.442 kmol/hr

Yield = 4488.62

4488.62+10.442x 100 = 99.77 %

3.6 Economic, Safety and Environmental Consideration

3.6.1 Economic

3.6.1.1 Desulphurization Unit

Following the technology evaluation, selection was done such that the optimum

performance is obtained in terms of cost and operation. Sulfatreat was chosen as sorbent

for desulphurization since it was among the cheapest costing $0.31/ lb (SulfaTreat,

2011). Moreover, no regeneration was carried out since it was found out that buying

new charge will be cheaper than sending for regeneration (Braga, 2004).

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3.6.1.2 Adiabatic Pre-reformer

Installation of an adiabatic pre-reformer (APR) reduces the reforming duty of

SMR as well as the fuel and energy consumption (Anonymous, n.d.). Similarly, if a

SMR of the same size is used, the plant throughput will increase about 5 – 10% with an

addition of APR unit (Anonymous, n.d.). The addition of an APR unit prolongs the

lifetime of downstream catalysts significantly (Munch et al., 2007). The overall

economical advantages of installing an APR are proven by reducing both capital and

operating cost.

3.6.1.3 Combined Reforming

The investment in the syngas generation accounts for 50% – 60% of the total

investment in a methanol production plant (Hansen and Nielsen, 2008). Natural gas

reforming is the cheapest and most efficient syngas generation technology as compared

to other feedstocks such as coal gasification and biomass (Hansen and Nielsen, 2008).

According to Haid and Koss (2001), conventional steam reforming is economically

applied to medium sized methanol plants and the maximum single train capacity is

limited to about 2500 mtpd. On the other hand, pure autothermal reforming (ATR) is

cheapest at capacities of 7000 mtpd (Hansen and Nielsen, 2008). However, it is found

that for mid-size capacities in the range of 2500 – 7000 mtpd, a hybrid two-step

(combined) reforming is the best choice as compared to conventional steam reforming

and pure autothermal reforming only (Hansen and Nielsen, 2008; Nielsen, 2008). That

is the one of the main reasons why in this design project, a two-step (combined)

reforming is chosen. A minimum of 3030 mtpd of methanol is produced. Therefore,

relative capital costs depend on capacity since the economy of scale is totally different

for steam reforming and autothermal reforming.

Besides that, an industrial study has been carried out to investigate the three

different synthesis gas technologies namely conventional steam reforming, combined

reforming and pure autothermal reforming (Hansen and Nielsen, 2008). Table 3.14

indicates the typical energy consumption and amount of circulating cooling water for

every tonne of methanol produced. Referring to Table 3.14, it is found that combined

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reforming requires the lowest energy (29.3 GJ/tonne) compared to steam reforming and

autothermal reforming. Also, this literature value reported by Hansen and Nielsen

(2008) is consistent with Cheng and Kung (1994). They state that the combined

reformer duty is reduced by 45% as compared to the base case of steam reforming.

However, the reduction in duty is mostly offset by the high cost of air separation unit

(ASU). As a result, on an overall basis, the combined reforming process which shows a

saving of 2.2% proves that this reforming configuration is the down-right choice for this

particular methanol plant (Cheng and Kung, 1994).

Table 3.14: Typical consumption value per tonne of methanol* (Hansen and Nielsen, 2008).

ParameterOne-step reforming

(SMR)Two-step reforming

(Combined)Pure ATR

Energy consumption (GJ/tonne)

31.0 29.3 31.0

Circulating cooling water (m3/tonne)

152 140 153

*Including drivers for oxygen plant, electricity and credit for steam export.

3.6.1.4 Methanol Synthesis

Although Boiling Water Reactor (BWR) could take up a high portion of the

capital cost, this reactor produces steam that could be used elsewhere in the process.

This greatly reduces the amount of steam to be purchased and thus reduces the operating

cost in the plant throughout the 25 years. This reactor also has a high yield and a low

thermal deactivation of catalyst due to its isothermal properties. A high yield of

methanol could generate more revenue whereas the low rate of catalyst deactivation

reduces the cost of buying and constantly replacing the catalyst. Therefore, the reactor

could be a profitable investment in the long run.

3.6.1.5 Methanol Purification

The high pressure separator is installed to separate the unreacted reactants from

the crude methanol. These reactants are then recycled to the reactor to increase the yield

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of methanol. Natural gas is sold at RM 600/t by Gas Malaysia Sdn. Bhd. Therefore, a

decrease in natural gas consumption will result in great savings in terms of operating

cost.

A separator is used as a substitute for the topping column in order to fit the

operation adequately due to the absence of various other volatile matters except for

methanol. Thus, additional cost involved in purchasing reboilers and condensers as well

as distillation columns can be reduced. Furthermore, distillation processes are

considered to be energy intensive processes. Also, since no cooling or heating is

required for the separator, the utility cost on water and steam consumption can be

reduced significantly.

In refining column, sieve plates are chosen seeing that the composition of the

methanol is relatively high in the feed. Simple perforated plates are sufficient to achieve

the desired purity of product. Sieve plate is the simplest and most economical type of

cross-flow plate as compared to others. Most importantly, sieve tray is effective over a

large range of flows with high capacities and does not foul easily. Hence, the

maintenance cost is reduced due to the ease of cleaning and high durability.

3.6.2 Safety Considerations


Much focus was also put in safety during the process synthesis. For

desulphurization, pressure and temperature were chosen properly for operation. Too low

pressure and too high pressure were avoided to prevent low reaction rates and high risk

of collapse as well as explosion.

3.6.2.2 Syngas Production

According to Cheng and Kung (1994), the methanol production using steam

reforming is a relatively clean and environmentally safe process.

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The process chosen in the conversion of syngas is a Low-Pressure Methanol

synthesis method where the process is set at a pressure and temperature of 90 bar and

220℃. This process is a safer process as compared to High-Pressure Methanol

synthesis where the process pressure temperature is 200-300 bar and temperature is 300-

400℃. This is due to the fact that at lowered pressure and temperature, the likelihood of

any serious explosion due to overpressure could be reduced.


The operation with high operating pressure has to be designed carefully in terms

of material selection and operating condition to avoid any failure of process or

equipment and thus frequent inspection and maintenance have to be provided. A second

separator with lower operating pressure is used instead of another higher pressure

separator to avoid higher chances of catastrophe happening from vessel explosion. Low

operating temperatures are preferred in the process to avoid the hazard leading to

equipment failure and unexpected disaster.

Methanol is classified as a primary class ‘Flammable liquids’ (Class 3.2) and

secondary class ‘Toxic substances’ (Class 6.1) (Methanex, 2010). In addition, methanol

vapor is considerably toxic to human which could cause visual disturbance, headaches,

dizziness, nausea and blurred vision. Owing to high flammability of methanol, the

operating temperature of distillation column is maintained at moderate temperatures

with pressure slightly higher than atmospheric pressure. Besides that, there is a

significant necessity to reduce the temperature of the liquid methanol product prior to

channeling it into the storage tank. Thus, additional condenser is placed to reduce

product temperature. The methanol produced must be properly stored in tightly closed

containers and in a well-ventilated area which is away from incompatible substances

such as heat sources and oxidizing agents (Microbial ID, 2009).

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3.6.3 Environmental Consideration


In the desulphurization section, one of the reasons for selecting sulfatreat is

because of its safe disposability. It is found out that in no instance has the spent media

absorbed enough material and subsequently released it upon undergoing tests such as

“hazardous metals test” or VOC test to be deemed hazardous waste (Braga, 2004).

3.6.3.2 Syngas Production

As natural gas is burnt to produce the heat required for the endothermic

reforming reaction, CO2 will be produced in the reformer furnace combustion zone. The

flue gas from the convection side of reformer contains CO2, and particulates. The

formation of NOx and VOCs are neglected since only negligible amounts are assumed to

be released. This is due to the assumption made that only complete combustion occurs

in the firebox of the steam reformer. Therefore, all VOCs are completely converted to

CO2 and H2O. On the other hand, NOx formation is neglected because according to

Smith (2005), thermal NO formation is negligible below 1300oC. Since the combustion

temperature in the radiant section was found to be 650oC, there would be no NOx

formation.

The main environmental objective is to reduce the CO2 emissions from the flue

gas. By reducing the CO2 emissions, the impact of methanol production on global

warming can be greatly reduced. Therefore, CO2 could be recovered from the flue gas

by applying pressure swing adsorption (PSA).


The formation of by-products in the BWR is low and therefore the discharge to

the environment poses less threat. The spent catalyst will be sent for regeneration

instead of disposing into environment which could reduce the environmental burdens.

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The separated stack gas is route to be burnt off and a certain amount of CO2 is

produced as flue gas. Through combustion, some of the components, which could result

in severe environmental impacts such as CH4 and CO, will be fully converted into CO2.

By this way, the impacts on the environment will be reduced and hence fulfill the

environmental regulations.

Wastewater generated from the distillation column comprises mainly water and

low concentration of acetic acid and methanol. Although concentration of acetic acid is

considerably low, it can contribute to aquatic toxicity. The pH value of acetic acid is

however, found to be lower than the allowable discharge limit of 5.5 to 9.0 according to

Malaysia industrial effluent discharge standard B (Water Treatment Resources, 2008).

On the other hand, there will be some degrees of methanol remained in the bottom of

the column due to its high solubility in water. The presence of methanol in wastewater

can cause adverse effect on aquatic life. The methanol content of wastewater should not

exceed 3.6 mg/L as suggested by the U.S. Environmental Protection Agent (EPA)

(Cheng and Kung, 1994). Hence, adequate wastewater treatment is essential to treat the

wastewater to the allowable discharge standards. However, single biological treatment

will be sufficient due to low concentrations of the contaminants present in the

wastewater stream (Methanex, 2010).

3.6.3.5 Summary of Contaminants

To summarize all the hazardous and non-hazardous contaminants discussed in

Sections 3.6.3.1 to 3.6.3.4, Table 3.15 shows the typical contaminants from various

sources in a methanol plant.

Table 3.15: Contaminants from Various Sources in a Methanol Plant.

Methanol Plant Effluents Contaminants

Flue gas in steam reformer CO2, particulates

Process condensate Total dissolved solids (TDS), total suspended solids (TSS)

Spent catalyst Various metals

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Cooling tower blowdown Total dissolved solids (TDS), total suspended solids (TSS)

Storage tank vent Methanol

Steam drum blowdown Total dissolved solids (TDS), total suspended solids (TSS)

Light ends Hydrocarbon (CH4), CO, CO2, CH3OH, CH3COOH

3.7 Process Optimization

3.7.1 Steam Reformer

The optimization for this section was done by manipulating the steam to carbon

ratio and observing the resulting composition of the effluent stream. Table 3.16 shows

the results of two different ratios used.

Table 3.16: Effluent composition using two different steam to carbon ratios.

Steam to Carbon Ratio 1.3 3.0

CH4 910.562 496.392

CO 937.172 1177.74

CO2 129.679 303.279

H2O 1602.99 4376.82

H2 3033.34 4449.45

N2 8.201 8.201

As shown above, by using a steam to carbon ratio of 3.0, more conversion of

methane is observed. Other than that, the production of CO and CO2 increases and since

both these components are reactants for the methanol synthesis process, the yield of

methanol can be increased as well. The main objective of a steam reformer in the

combined reforming configuration is to produce a H2-rich syngas. This is also achieved

by increasing the ratio to 3.0. Many other ratios were attempted as well. Collectively, a

steam to carbon ratio of 3.0 was chosen. Although the benefits mentioned above are at

the expense of high steam requirements, the steam utilized here is part of the recycled

steam produced within the plant.

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3.7.2 Autothermal Reforming

In addition to that, optimisation was also carried out in autothermal reforming.

Seeing as there is a wide range of O2 to carbon ratio ranging from 0.6 to 1.5 (Logdberg

and Jakobsen, 2010; Zamaniyan et al., 2009), this ratio could be adjusted so that a

stoichiometric ratio of close to 2.0 in the syngas was attained. This ratio is one of the

main requirements in producing a good yield of methanol in the methanol synthesis

process as reported by most of the literatures (Logdberg and Jakobsen, 2010; Lurgi,

2006; Hansen and Nielsen, 2008; Petersen et al., 2004; Uhde, 2006; Van Den

Oosterkamp and Van Den Brink, 2010). Table 3.17 shows the effect of O2 to carbon

ratio on the stoichiometric ratio of synthesis gas. After optimisation, it clearly indicates

that O2 to carbon ratio of 0.8 is the best choice as compared to other ratios. The ratio of

0.6 is not chosen since a low conversion of 20.5% of CH4 is obtained in the catalytic

zone.

Table 3.17: Effect of O2 to carbon ratio on stoichiometric ratio of syngas

O2 to carbon ratio Stoichiometric ratio, SR

0.6 2.10

0.8 1.85

1.0 1.61

1.2 1.37

1.5 1.00

3.7.3 Methanol Synthesis

The production capacity without the recycle stream in the process is clearly

lower than that with the recycle stream. A scale up in the flow rate of natural gas feed is

required in order to obtain the same methanol production capacity as the system with

recycle stream. This would therefore increase the feedstock usage as well as incur a

higher cost. Table 3.18 shows the comparison between a process with a recycle and

without a recycle stream.

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Table 3.18: Effect of recycle and without recycle process on the conversion and selectivity

ComponentInlet to reactor

(kmol/hr)(With Recycle)

Outlet from reactor (kmol/hr)

(With Recycle)

Inlet to reactor (kmol/hr)

(Without Recycle)

Outlet from reactor (kmol/hr)(Without Recycle)

CH4 359.24 359.24 47.24 47.24

CO 4112.1 183.41 3952.76 73.33

H2O 1121.37 1748.0 1121.37 1443.36

CO2 2572.7 1946.1 882.51 560.52

H2 10702.87 986.92 9845.73 1131.26

N2 153.99 153.99 20.25 20.25

CH3OH 0.00 4534.1 0.00 4191.07

CH3COOH 0.00 10.55 0.00 5.17

Conversion 95.36% 98.14%

Selectivity 99.78% 99.88%

3.7.4 Methanol Purification

Due to inappropriate operation for distillation column, a separator is used as

substitute for separation of light-end gas from methanol product. By adopting this

technology, a significant cost saving can be achieved due to the absence of distillation

column which has a great utility consumption and is relatively expensive due to its

complexity of construction and additional reboiler and condenser required. With a

relatively cheaper separation being used in this process, less capital investment can be

achieved by satisfying the product quality requirment.

On the other hand, the separated gas will be utilized instead of releasing directly

into the atmosphere which might cause adverse effects to the environment. The

recycling of gas to the methanol synthesis can increase the product yield with less raw

material consumption in the process. The remaining gas from purification process will

be routed to burner to convert most of the unreacted substances into CO 2 and captured

through PSA unit. The inert compounds will be released through stack tower while the

captured CO2 will be stored through carbon capture and sequestration (CCS) system and

further supplied as valuable industrial gas to the chemical industry such as refrigeration

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SYNTHANOL SDN. BHD.


systems, inert agent for food packaging and many other applications (Mazzotti, n.d.).

This technology will significantly reduce the environmental burden and hence increase

the process sustainability with cleaner process.

3.8 Process Flow Diagram

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3.9 Process Flow with Reference to Process Flow Diagram

Feedstream 1.1, the natural gas feedstock, available at 30°C and 30 bar, is

preheated in the convection section of the steam-methane reformer, R-104 to 220°C and

compressed to 40 bar by compressor C-101. The preheated stream 1.3 is fed to the

desulphuriser unit R-101. The exit from this first unit goes into a second desulphuriser

unit R-102 for further purification. The desulphuriser operates adiabatically at a

pressure of 40 bar and a temperature between 235 °Cand 239 °C. The desulphuriser exit

stream 1.4, at 235.58 °C, is fed to the saturator V-101. Stream 1.6, is the make-up water

for the saturator entering as subcooled liquid at 80°C and 38 bar. This is mixed with the

liquid effluent from the saturator, 1.5 and enters as recycle stream 1.7 to the saturator.

Pump P-101 pumps the liquid effluent back to the saturator operating pressure of 38 bar.

The exit stream from the saturator is stream 1.8.

After exiting from the saturator, V-101, the process stream 1.8 flows into the

convection section of the steam reformer, being preheated to 538.29°C by the flue gas.

This preheated stream 2.1 is flowing at 38 bar. It is regulated to 36 bar, stream 2.2, and

enters the pre-reformer. At the same time, the compressed steam, 2.3, enters at 36 bar.

After pre-reformer, the effluent stream, 2.4 will exit at 500°C and 35.6 bar, with a

pressure drop of 0.4 bar. This stream diverges into streams 2.5 and 3.1, which will enter

both ATR and SMR respectively.

The effluent from APR which is process stream 3.1 flows at 500 °C and 35.6

bar. This stream then combines with preheated steam 3.3 at the same conditions. The

preheated steam is produced by waste heat boiler, E-102. The resultant process stream

3.4 enters the steam reformer, R-105 at 500 °C and 35.6 bar. Simultaneously, air stream

3.6 at 330 °C and 1.6 bars, natural gas (fuel) stream 3.2 at 30 °C and 30 bar as well as

recycle stream 6.11 at 40 °C and 10 bar all enter the firebox, R-104 to be combusted.

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The air stream 3.6 is initially preheated in the convection section of the steam reformer

R-105, from 30°C and 1.013 bar to 330 °C and 1.6 bars prior to being combusted.

The flue gas stream 3.8 exits the steam reformer at 60.9 °C and 1.26 bar and is

cooled to 40 °C and 1.24 bar in stream 3.9. This stream then enters the pressure swing

absorption (PSA) vessels, R-109 and R-110 in an alternative manner. The absorbed CO2

is released in stream 3.12 at 40 °C and 2.03 bar whereas the PSA effluent 3.10 at 40 °C

and 30.4 bar is blown through a flue gas fan, F-102 to stack at same conditions.

After exiting from the SMR, R-105, the effluent stream 3.7 will be cooled from

850 oC to 381.3 oC. The cooled effluent gas will combine with the bypassed APR stream

4.1, oxygen stream, 4.3 and steam stream, 4.4. The supply oxygen (stream 8.9) is

preheated from 30 oC and 30 bar to 230 oC and 30 bar in the convection section of the

steam reformer, R-104. The pressure of all the feeds entering the ATR is around 30 bar.

After ATR, the effluent stream 4.5 will leave at 1000 oC and 29 bar with a pressure drop

of 1 bar. The hot reformed gas (stream 4.5) will enter a heat exchanger, E-102 to cool

the main process stream to 952.2 oC and 29 bar. Meanwhile, saturated steam of 244.2 oC

and 36 bar(stream 8.8) is superheated to a higher temperature of 500 oC (stream 3.3)

which will enter the SMR together with the effluent from APR (stream 3.1). After that,

the effluent stream 4.6 will be further cooled down to 300 oC and 29 bar in a waste heat

boiler, E-103 using cooling water medium (stream 8.16). Here the cooling water of 30 oC (stream 8.16) is superheated to 250 oCand 15 bar (stream 8.17). The superheated

steam (stream 8.17) will enter a steam turbine, T-101 where electricity is generated

which could be used within the industry process. Saturated steam (streams 8.18) leaves

the turbine at 179.9 oC and 10 bar. Part of the saturated steam (stream 8.20) will be

recycled and compressed together with the supply steam (8.2) before entering each

reformer reactor. The remaining saturated steam (stream 8.19) will be sent to the

reboiler, E-109 in the distillation column, D-101. The effluent (stream 4.7) from the

waste heat boiler, E-103 will have to be reduced from a temperature of 300 oC to 80 oC

(stream 4.8) in order to condensate and remove the water from the main process stream

in a knock out drum, S-101. In the knock out drum, S-101, all the non-condensable

gases will leave the separator as vapor phase (stream 4.9) whereas the bottom

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SYNTHANOL SDN. BHD.


condensate (stream 4.12) will be pumped to 80 oC and 38 bar (stream 4.13) to the

saturator, V-101 and cooling tower, E-110.

The separated syngas of a stoichiometric ratio of 1.85 (stream 4.9) from the

reforming section with a temperature of 80oC and a pressure of 29 bar is compressed

using compressor C-102 to 148.3oC and 94 bar (stream 4.10). This stream will mix with

stream 6.7 and then preheat using E-105 to 220℃ (stream 5.1) before entering into two

methanol converters in equimolar ratio (stream 5.2 and stream 5.3). These streams then

enter the two boiling water reactors (BWR) (Reactor R-107 and R-108) respectively at a

pressure of 94 bar. The exothermic heat from the reaction is used to heat the boiling

water in the reactor and produce a mixture of saturated steam and water (stream 5.11

and stream 5.12) before entering a steam drum, V-102 to be separated and produce

saturated steam (stream 5.14) at 10 bar and 179.9℃ which is to be fed to the reboiler,

E-109. Feed water (stream 5.7) is fed at 30℃ and 10 bar to replenish the water in the

drum that is converted to steam and leave the steam drum. The removal of exothermic

heat is used to maintain the isothermal conditions in the reactor. The conversion of

carbon monoxide to methanol is approximately 95%. The outlet of the two reactors

(stream 5.4 and stream 5.5) containing 2% CO, 20% CO2, 10.7% H2 at 220℃ and 90

bar will be condensed to 40℃ and separated in a high pressure separator, S-102 and

96.5% of the resulting overhead product containing most of the gases will be

compressed to 94 bar using compressor C-103 and recycled back to the process stream

before re-entering the reactors (R-107 and R-108) whereas the remaining gas will be

regulated from 90 bar to 10 bar. The bottom product (stream 6.2) from the separator, S-

102 is passed through a letdown valve, V-13 to reduce the pressure to 10 bar before

entering a letdown vessel, S-103 where liquid is flashed and flowed upwards with gases

(stream 6.10). The installation of mist eliminator can significantly retain 99% of liquid

in the bottoms product while all the remaining gases will be separated as overhead

product (stream 6.10) and mixed with the purge gas (stream 6.5) from high pressure

separator, S-102 to be burnt in the reformer burner, R-104. The retained liquid product

will be separated in the bottom (stream 6.8) and pumped to 5 bar into a distillation

column for further purification.

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Stream 6.9, the effluent from letdown vessel, S-103 comprises liquid methanol,

acetic acid and water at temperature of 40℃ and pressure of 10 bar. This is pumped

into distillation column, D-101 for further purification in order to obtain methanol

product of 99.85% purity. Feed stream 7.1 enters at 40℃ and 5 bar. Then the feed will

flow down the distillation column, D-101 through the sieve trays and stream 7.9 enters

the reboiler, E-109 at 90℃and 2.30 bar. Vaporization of methanol occurs in the reboiler

where the methanol vapor is channeled back into the column at temperature of 124.5℃

and pressure of 2.30 bar as shown as stream 7.1. The remaining water, small

concentration of acetic acid and non-vaporized methanol exit as wastewater stream 7.11

at the same temperature and pressure as stream 7.1.

Stream 7.2, the top product of the distillation column is methanol vapor which

exits at a temperature of 85℃ and pressure of 1.87 bar. All methanol vapor is

condensed at the distillation column condenser, E-107 and condensed liquid stream 7.3

is stored in reflux drum, V-103 at a temperature of 83.32℃ and pressure of 1.87 bar. A

fraction of the condensed methanol is sent back into the refining column, D-101 and the

remaining is directed into storage tank as dictated by the reflux ratio illustrated by

streams 7.7 and 7.5 respectively. After methanol product cooler, E-108, the product

stream 7.6 is at temperature of 45℃ and pressure of 1.87 bar. The product is then stored

at storage tank, V-104.

Saturated steam from steam drum, V-102 is fed into the reboiler at a temperature

of 179.9℃ and pressure of 10 bar. The saturated steam is then cooled to 130℃ and 10

bar in stream 7.13. Cooling water stream 7.14 originated from cooling tower, E-109 at

35℃ and 10 bar is used to condense the methanol product in the distillation column

condenser E-107. The cooling water stream 7.15 leaves the distillation condenser at 77

℃ and 10 bar. Cooling water stream 7.16 of temperature 35℃ and pressure 1 bar is

used to cool methanol product in E-108 . Similarly, the cooling water stream 7.17 exits

the distillation condenser at 40℃ and 1 bar. Both cooling water will then be recycled

back into cooling tower, E-110.

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3.10 Energy Integration

3.10.1 Heat Exchanger Network (HEN) Design

There are two different methodologies developed for heat integration in

chemical processes namely Heat Exchanger Network Synthesis (HENs) and Pinch

Technology (Martin et al., 2008).

Heat Exchanger Network Synthesis (HENs) method is generally solved with

software programmed based on mixed integer non-linear (MINLP) optimization of

superstructures of possible exchanger options (Martin et al., 2008). This tool is able to

establish the best solutions for HENs problem. The second methodology is known as

Pinch Technology. The advantage of using pinch technology is the ability to optimize

the number of heat exchangers, heat exchanger area and minimize capital, production as

well as utility costs using present energy stream with high or low energy content

(Klemes et al., 2011).

Pinch technology identifies the heat sources (hot streams) and heat sinks (cold

streams) from the process flow and represents it on temperature-enthalpy diagram

(Klemes et al., 2011). The position of “pinch” is determined by the graphical

representation in the form of composite curves with the incorporation of minimum

temperature for heat exchange. It usually occurs between the hot and cold streams curve

where the region above the pinch is the heat sink and below the pinch is the heat source.

Heat integration was performed on the following sections of the methanol plant.

Firstly, the streams, which require heat recovery, were identified. Six cold streams and

four hot streams were integrated and tabulated in Table 3.19. Referring to the PFD, the

cold streams, which require heating, include streams 1.1, 1.8, 8.9, 3.5, 4.11 and external

cooling water utility, namely natural gas feed before entering into desulphuriser units,

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saturated natural gas leaving the saturator, supply oxygen feed, supply air and methanol

converter feed. On the other hand, the hot streams, which require cooling, encompass

streams 3.7, 4.7, 8.4 and 3.8, namely ATR feed from SMR effluent, ATR effluent,

cooling of compressed steam at constant pressure and flue gas exiting the convection

side of the firebox. A minimum temperature difference between hot stream and cold

stream, ∆ T min of 10oC is chosen.

After that, a problem table algorithm (as shown in Table 3.20) is built based on

the stream populations. Then a composite curve (as shown in Figure 3.30) is plotted

based on the temperature and enthalpy for each stream. Finally, a heat exchanger

network (HEN) of all the nine streams is simulated using Aspen Energy Analyzer

Version 7.2. The network is shown in Figure 3.34.

Table 3.19: Stream table for hot and cold process streams.

Stream Type ṁ (kg/h) Cp (MW/ᵒC)1.1 Cold 30 220 2.11 80216.1 0.04702 35 2251.8 Cold 279.66 538.29 2.643 144868.9 0.10636 284.66 543.293.5 Cold 30 330 1.011 275143.1 0.07727 35 3353.7 Hot 850 381.3 2.95 157980.2 0.12946 845 376.33.8 Hot 60.9 40 1.107 301750.6 0.09279 55.9 354.7 Hot 300 80 2.36 372368.5 0.24411 295 75

4.11 Cold 139.1 220 2.2 279986.5 0.17110 144.1 2258.4 Hot 358.4 244.2 2.315 150547.6 0.09681 353.4 239.28.9 Cold 30 230 0.9132 76952.1 0.01952 35 235

Ts (ᵒC) TT (ᵒC) Cp (kJ/kg.K) Ts*(ᵒC) TT*(ᵒC)

Ts* and TT* are the shifted temperatures for supply and target temperatures of the process stream.

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Table 3.20: Problem Table Algorithm.

845301.71 -0.12946 -39.0582 S

543.29166.99 -0.02310 -3.8572 S

376.322.9 0.10636 2.4356 D

353.418.4 0.00955 0.1757 D

33540 0.08682 3.4727 D

29510.34 -0.15729 -1.6264 S

284.66

45.46 -0.26365 -11.9855 S

239.24.2 -0.16684 -0.7007 S

23510 -0.18636 -1.8636 S

22510 0.07080 0.7080 D

21570.9 0.07080 5.0197 D

144.149.1 0.07080 3.4763 D

9520 0.55969 11.1939 D

7519.1 0.80380 15.3526 D

55.920.9 0.71101 14.8602 D

35

Temperature interval (ᵒC)

ΔTinterval

(ᵒC)∑Cpc - ∑CpH

(MW/K)ΔHinterval (MW)

Surplus/Deficit

1.1 3.5 8.9 CW

1.8

3.7

4.7

8.4

4.11

3.8

Cp

= 0

.04702

Cp

= 0

.10636

Cp

= 0

.07727

Cp

= 0

.0927

9

Cp

= 0

.24411

Cp

= 0

.1711

0

Cp

= 0

.01952

Cp

= 0

.6600

Cp

= 0

.09681

Cp

= 0

.12946

In order to design the heat exchanger network from the Apen Energy Analyzer,

the supply and target temperatures of each stream are inserted as shown in Table 3.21.

Figures 3.31 and 3.33 summarise the details of each heat exchanger as well as the

network cost and performance. It is found that for all ten heat exchangers, a total area of

9339 m2 and a total number of shells of 27 are required. This configuration is chosen

since it provides the smallest area and number of shells as compared to the preliminary

design configuration which is summarized in Figure 3.32. Due to the extremely large

cross sectional area (4.155×106 m2), large number of shells (8327) and high heat load

(1.385×1011 kJ/hr) for the preliminary heat exchanger design as indicated in Figure 3.32,

this design configuration is not chosen. Both the preliminary and final heat exchanger

network (HEN) design are summarised in Table 3.22.

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Table 3.21: Supply and Target Temperatures of Each Stream from Aspen Energy Analyzer.

Figure 3.30: Composite curves from Aspen Energy Analyzer.

Figure 3.31: Summary of all heat exchanger details from Aspen Energy Analyzer

(Chosen).

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Figure 3.32: Summary of Preliminary Heat Exchanger Network Design from Aspen Energy Analyzer (Not Chosen).

Table 3.22: Summary of Parameters for Preliminary and Final Heat Exchanger Network (HEN) Design.

ParametersPreliminary HEN Design

(Not chosen)Final HEN Design

(Chosen)

Cost Index 9.596×108 2.410×106

Area (m2) 4.155×106 9339

Number of shells 8327 27

Heat Load (kJ/hr) 1.385×1011 4.585×108

Figure 3.33: Summary of Overall Heat Exchanger Network Cost and Performance from Aspen Energy Analyzer.

In overall, this design network is chosen to transfer the heat from the hot streams

to the cold streams as much as possible without using any external utilities. By this way,

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a maximum heat recovery and a minimum consumption and cost of the hot and cold

utilities would be achieved. This is because the capital cost is dependent on the number

of exchangers in the network as well the size (area).

A very important aspect of heat integration is the capital and total costs incurred.

Heat integration synthesises heat exchanger networks to keep the costs at minimum.

This is mainly done by optimising the major components of the heat exchange networks

contributing to the capital costs. These include the number of units, the heat exchange

area, the number of shells, the material of construction, the heat exchanger type as well

as the pressure rating (Smith, 2005).

The number of units refers to the number of matched between the hot and cold

streams. Generally, to get minimal capital cost, the final heat exchanger network uses a

minimum number of units. This is usually achieved by having zero independent loops in

the network and maximum number of components (Smith, 2005). However, the safest

assumption for the number of components is one such that for a loop free network, the

minimum number of units is given y the number of streams minus one.

A minimum heat exchange area also contributes to achieving a lower capital

cost. Information to predict the minimum network area is obtained from balanced

composite curves which have no residual demand for utilities. The minimum area can

then be calculated given the overall heat transfer coefficient and the log mean

temperature difference (Smith, 2005).

Another requirement to target minimal capital cost is to have the least possible

number of shells. Countercurrent devices use a number of shells equal to the number of

units. But, usually a balance is kept to maintain consistency between achieving

maximum energy recovery and the corresponding minimum number of units target.

The other factors, material of construction, type of exchanger and pressure

rating, all affect the capital cost of a single heat exchanger with surface area A

according to the following relationship (Smith, 2005):

Installed cost = a+b Ac

where a, b, c are the cost law constants that incorporate the aforementioned factors.

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The material of construction, in particular, is chosen according to the heat load and

nature of the stream

The final heat exchange network design produced from heat integration can then

be considered to be the optimum design that will achieve the required performance at

the lowest cost efficiently. This basically underlines the importance of carrying out heat

integration in this methanol plant.

The process flow diagram after performing heat integration is as shown in

Section 3.10.2.

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Figure 3.34: Chosen Heat Exchanger Network (HEN) Design from Aspen Energy Analyzer.

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3.10.2 Process Flow Diagram With Heat Integration

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