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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
[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).
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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)
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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
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
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.
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Baer, B. (2011), The Types of Distillation Trays (Online). Retrieved on 5th of August 2011,
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Reactions’. In: Fundamental of Industrial Catalytic Process (2nd Edition), New Jersey:
Wiley-Interscience, pp. 382 – 398.
Bhandari, D.A., Bessho, N. and Koros, W.J. (2010), ‘Hollow Fiber Sorbents for
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