submission a report (version4)
TRANSCRIPT
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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SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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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SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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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SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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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SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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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SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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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SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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).
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SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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).
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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%.
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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
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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)
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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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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).
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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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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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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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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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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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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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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%
(Logdberg and Jakobsen, 2010).
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).
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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:
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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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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
(Logdberg and Jakobsen, 2010).
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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
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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).
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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
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(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
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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).
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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.
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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).
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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).
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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)
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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)
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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).
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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
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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.
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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).
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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
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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).
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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,
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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.
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Figure 2.16: ICI quench reactor (GBH Enterprsise, n.d.).
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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
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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.
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Figure 2.17: Spherical reactor in series (Cheng et al., 1994).
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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.
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Figure 2.18: Tube cooled reactor design (GBH Enterprsise, n.d.).
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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).
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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.
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Table 2.9: Review of different catalyst for methanol synthesis (Mäyrä et al.,2008).
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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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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
3.6.2.1 Desulphurization Unit
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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3.6.2.3 Methanol Synthesis
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.
3.6.2.4 Methanol Purification
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
3.6.3.1 Desulphurization Unit
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).
3.6.3.3 Methanol Synthesis
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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3.6.3.4 Methanol Purification
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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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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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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SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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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SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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,
107
SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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.
108
SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
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.
109
SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
Figure 3.34: Chosen Heat Exchanger Network (HEN) Design from Aspen Energy Analyzer.
110
SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
3.10.2 Process Flow Diagram With Heat Integration
111
SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
112
SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
113
SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
114
SYNTHANOL SDN. BHD.
SUBMISSION A FOR DESIGN OF METHANOL PRODUCTION PLANT IN KUANTAN
115
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