revisi tk11 report assignment 1

UNIVERSITAS INDONESIA

PRELIMINARY DESIGN OF ETHYLENE PRODUCTION FROM COAL

GASIFICATON

REPORT ASSIGNMENT 1

GROUP 11

GROUP PERSONNEL:

Ikhsan Nur Rosid (1106007691)

Ikhwan Muttaqin (1106010925)

Nuri Liswanti Pertiwi (1106015421)

Ranti Fabrianne (1106020522)

Sirly Eka Nur Intan (1106005055)

CHEMICAL ENGINEERING DEPARTMENT

ENGINEERING FACULTY

UNIVERSITAS INDONESIA

DEPOK

2014

EXECUTIVE SUMMARY

Entering the era of free trade, Indonesia is required to be able to compete with other countries in the field of industry. Ethylene is one of the important compounds in the petrochemical industry chain and organic chemicals largest in the world. Until 2005, the only plant in Indonesia which produces ethylene is PT. Chandra Asri Petrochemical Indonesia. Ethylene products from PT. Chandra Asri is almost all of those consumed of polymer grade, which is mostly used as a raw material of Linear Low Density Poliethyelene (LLPDE) Poliethylene Plant and High Density (HDPE) Plant PT. Chandra Asri, while a small portion is sold to PT. Peni and PT. Asahimas Subentra Chemical. Considering the needs for ethylene increasing both ethylene groups of chemical grade and polymer grade, while ethylene producers themselves can be said is still limited, it can be said that the market share for the ethylene plant is still very open, both domestic and foreign markets.

The raw material of ethylene plant is coal. In Indonesia, the condition of coal reserves is very abundant and is the energy source with the largest reserves compared with natural gas and petroleum. Considering the carbon composition, price, and hardness then bituminous coals are selected. While coal suppliers selected from PT Bukit Asam and PT Adaro which is located in southern Sumatra. The supply of oxygen can be fulfilled by some companies such as PT. Air Liquid and PT. Air Product. We preferred to PT. Air Product Indonesia, since they has technology which called oxygen production on-site and mostly produce gaseous oxygen. By 2014, the MSE is Rp2.443.000 in Cilegon. That number are considered rational for the petrochemical industry because of the labor in this sector is mostly skilled and trained workers.

Each gasifier has advantage and disadvantage, such as fixed bed has advantage in high conversion and total of gasification media is small, fluidized bed has advantage in feedstock (coal rank) because all variant of coal capable become feed, and entrained bed has advantage in high purity of product. From scoring, we choose gasifier that will be used is Entrained Bed Reactor. This type of gasifier is choosen because high purity of product. Meanwhile, ethylene synthesis is done using Methanol-to-Olefin technology based on scoring. This plant is then simulated via ASPEN HYSYS 7.3. The mass efficiency of this plant is 45,47% while the energy efficiency is 27.080 kJ/kg.

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CONTENTS

EXSECUTIVE SUMMARY.................................................................................ii

CONTENTS..........................................................................................................iii

LIST OF FIGURE.................................................................................................v

LIST OF TABLE..................................................................................................vi

CHAPTER I INTRODUCTION...........................................................................1

1.1. Background...............................................................................................1

1.2. Review Literature......................................................................................2

1.1.1. Ethylene.............................................................................................2

1.1.2. Coal....................................................................................................5

1.1.3. Types of Coal and Its Application.....................................................6

1.1.4. Coal Gasification................................................................................7

1.2. Market Analysis........................................................................................8

1.3. Capacity Analysis....................................................................................10

1.4. Raw Material Analysis............................................................................12

1.5. Plant Location Analysis...........................................................................16

CHAPTER II PROCESS SELECTION............................................................20

2.1. General Process...........................................................................................20

2.2. Alternative Process..................................................................................20

2.2.1. Gasification Technology..................................................................21

2.2.2. Acid Gas Removal Technology............................................................21

2.2.4. Synthesis Ethylene................................................................................22

2.3. Process Selection.....................................................................................28

2.3.1. Gasification Technology..................................................................28

2.3.2. Synthesis Ethylene...........................................................................35

2.4. Process Description.................................................................................36

2.4.1. Block Flow Diagram........................................................................36

2.4.2. Process Flow Diagram.....................................................................38

2.4.3. Process Description..........................................................................43

CHAPTER III MASS & ENERGY BALANCE...............................................57

3.1. Mass Balance for Equipment..................................................................57

3.1.1. Gasification Unit..............................................................................57

3.1.2. Acid Gas Removal Unit...................................................................60

3.1.3. Water Gas Shift Unit........................................................................64

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3.1.4. Synthesis Methanol Unit..................................................................68

3.1.5. Synthesis Ethylene Unit...................................................................71

3.2. Energy Balance for Equipment...............................................................76

3.2.1. Gasification Unit..............................................................................76

3.2.2. Acid Gas Removal Unit...................................................................76

3.2.3. Water Gas Shift Unit........................................................................77

3.2.4. Synthesis Methanol Unit..................................................................78

3.2.5. Synthesis Ethylene Unit...................................................................79

3.3. Overall Mass Balance..............................................................................80

3.4. Overall Energy Balance...........................................................................80

3.5. Mass Efficiency.......................................................................................81

3.6. Energy Required per Unit Product and Energy Efficiency.....................81

CHAPTER IV CONCLUTION..........................................................................82

REFERENCES.....................................................................................................83

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LIST OF FIGURE

Figure 1.1. Coal and Utilization........................................................................................6

Figure 1.2. Coal Gasification Process...............................................................................7

Figure 1.3. Prediction Cpacity Production, Consumption, Imports..................................10

Figure 1.4. Coal Development Process.............................................................................13

Figure 1.5. Coal Quality Based on Coal Type..................................................................13

Figure 1.6. Plant Location.................................................................................................17

Figure 1.7. Indonesia’s Coal Reverse...............................................................................17

Figure 1.8. Target Market, Coal Mining, and Plant Location...........................................18

Figure 2.1. Black Box of This Plant..................................................................................20

Figure 2.2. Simple Flow Diagram for the MTO process by UOP....................................25

Figure 2.3. General Block Diagram Flow of Fischer Tropsch..........................................27

Figure 2.4. Fixed Bed Gasifier..........................................................................................31

Figure 2.5. Fluidized Bed Gasifier....................................................................................32

Figure 2.6. Entrained Bed Gasifier...................................................................................33

Figure 2.7. Block Flow Diagram for This Plant...............................................................37

Figure 2.8. Process Flow Diagram for Gasification Unit.................................................38

Figure 2.9. Process Flow Diagram for Acid Gas Removal Unit.......................................39

Figure 2.10. Process Flow Diagram for Water Gas Shift Unit.........................................40

Figure 2.11. Process Flow Diagram for Synthesis Methanol Unit...................................41

Figure 2.12. Process Flow Diagram for Synthesis Ethylene Unit....................................42

Figure 2.13. Methanol to Olefins Reactor.........................................................................51

Figure 2.14. Reaction to Produce Ethylene and Propylene...............................................51

Figure 2.15. Diagram of Temperature versus Selectivity.................................................52

Figure 2.16. Feed Effect to Selectivity.............................................................................53

Figure 3.1. The Hysis Simulation for Gasification Unit...................................................57

Figure 3.2. The Hysis Simulation for Acid Gas Removal Unit........................................60

Figure 3.3. The Hysis Simulation for Water Gas Shift Unit.............................................64

Figure 3.4. The Hysis Simulation for Synthesis Methanol Unit.......................................68

Figure 3.5. The Hysis Simulation for Synthesis Ethylene Unit........................................72

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LIST OF TABLE

Table 1.1. Data Capacity Production, Consumption, Exports, Imports Ethylene Year

Periode 2005 - 2009......................................................................................9

Table 1.2. Average Coal Price According to the Type.....................................................14

Table 1.3. Raw Material Selection....................................................................................14

Table 1.4. The Composition of Bituminous Coal.............................................................14

Table 1.5. Bituminous Coal Physical Properties...............................................................15

Table 1.6. Oxygen Physical Properties.............................................................................16

Table 2.1. Scoring the type of Gasifier.............................................................................34

Table 2.2. Scoring Synthesis Ethylene..............................................................................35

Table 2.3. Condition Operation in Each Equipment (Methanol to Olefin).......................55

Table 2.4. Condition Operation in Refrigeration Train....................................................55

Table 3.1. Mass Balance on Mixer...................................................................................57

Table 3.2. Mass Balance on Slurry Pump.........................................................................58

Table 3.3. Mass Balance on Gasifier................................................................................58

Table 3.4. Mass Balance on Cooler Raw Syngas.............................................................59

Table 3.5. Mass Balance on Cyclone Separator................................................................59

Table 3.6. Mass Balance on Raw Syngas Compressor.....................................................60

Table 3.7. Mass Balance on Raw Syngas Cooler.............................................................61

Table 3.8. Mass Balance on Absorber Column................................................................61

Table 3.9. Mass Balance on Rich DEPG Pump................................................................62

Table 3.10. Mass Balance on Stripper Column................................................................62

Table 3.11. Mass Balance on Rich DEPG Cooler............................................................63

Table 3.12. Mass Balance on Rich DEPG Pump..............................................................63

Table 3.13. Mass Balance on Expander............................................................................64

Table 3.14. Mass Balance on Reboiler.............................................................................65

Table 3.15. Mass Balance on High Temperature Shift.....................................................65

Table 3.16. Mass Balance on Reboiler.............................................................................66

Table 3.17. Mass Balance on Low Temperature Shift......................................................67

Table 3.18. Mass Balance on Cooler................................................................................67

Table 3.19. Mass Balance on Compressor........................................................................69

Table 3.20. Mass Balance on Heat Exchanger..................................................................69

Table 3.21. Mass Balance on Reactor...............................................................................70

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Table 3.22. Mass Balance on Separator............................................................................70

Table 3.23. Mass Balance on Tank...................................................................................71

Table 3.24. Mass Balance on Distillation Column...........................................................71

Table 3.25. Mass Balance on Distillation Column...........................................................72

Table 3.26. Mass Balance on MTO Reactor.....................................................................73

Table 3.27. Mass Balance on Water Separator.................................................................73

Table 3.28. Mass Balance on Demethanizer.....................................................................74

Table 3.29. Mass Balance on Deethanizer........................................................................75

Table 3.30. Mass Balance on Ethylene Tower..................................................................75

Table 3.31. Gasification Unit Energy Balance.................................................................76

Table 3.32. Acid Gas Removal Unit Energy Balance.......................................................77

Table 3.33. Water Gas Shift Unit Energy Balance...........................................................78

Table 3.34. Methanol Synthesis Energy Balance.............................................................79

Table 3.35. Energy Balance on Synthesis Ethylene..........................................................79

Table 3.36. Mass Balance Overall....................................................................................80

Table 3.37. Energy Balance Overall.................................................................................81

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CHAPTER IINTRODUCTION

1.1. BackgroundEntering the era of free trade, Indonesia is required to be able to compete

with other countries in the field of industry. Industrial development in Indonesia

are very influential on the resistance of the Indonesian economy. Many sectors of

the chemical industry are holding role to the improving industry in Indonesia. The

innovation of production processes and new plant construction oriented to

reducing our dependence on foreign products and to increase foreign exchange is

required, one of which with the addition of ethylene plant.

Ethylene is one of the important compounds in the petrochemical industry

chain and organic chemicals largest in the world. Ethylene is the basic ingredient

for various intermediate products and final products such as plastics, resins, fibers,

elastomers, solvents, surfactants, coatings, and antifreeze. Generally, ethylene

products are divided into two groups, namely polymer grade and chemical grade.

Polymer grade has a purity of up to 99%, while for chemical grade purity ranged

from 92 until 94%. Polymer grade is the largest consumer of the raw material of

ethylene, which is up to 45% of the total production of ethylene. Besides the

polymer grade, ethylene was also consumed by a group of chemical grade. These

groups include chemical grade such as ethanol, ethylene oxide, vinyl acetate,

ethylene solvents, and so on.

Until 2005, the only plant in Indonesia which producing ethylene is PT.

Chandra Asri Petrochemical Indonesia. Ethylene products from PT. Chandra Asri

is almost all of those consumed of polymer grade, which is mostly used as a raw

material of Linear Low Density Poliethyelene (LLPDE) Poliethylene Plant and

High Density (HDPE) Plant PT. Chandra Asri, while a small portion is sold to PT.

Peni and PT. Asahimas Subentra Chemical. Along with the development of the

times, the needs for ethylene is also increased, up to the year 2009 there were

already four plant in Indonesia that produces ethylene, namely PT. Salim Group,

PT. Pertamina, PT. Shell, and PT. Chandra Asri Petrochemical.

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The development of domestic ethylene import according BPPN statistics,

ethylene imports in 2010 reached 641.0000 tons. The needs for polymer grade of

ethylene in the country in 2010 was 1.3 million tons per year. Global ethylene

demand in 2015 is expected to reach 160 million tons, which each year will

experience an increase of 5.5%. Considering the needs for ethylene increasing

both ethylene groups of chemical grade and polymer grade, while ethylene

producers themselves can be said is still limited, it can be said that the market

share for the ethylene plant is still very open, both domestic and foreign markets.

Product ethylene produced can be sold for the benefit of polymer grade or

chemical grade.

The raw material of ethylene plant is coal. In Indonesia, the condition of

coal reserves is very abundant and is the energy source with the largest reserves

compared with natural gas and petroleum. Coal reserves are spread almost each

island in Indonesia, with the biggest reserves are in the island of Sumatra and

Balikpapan. Based on data from the Geological Agency, Ministry of Energy and

Mineral Resources, Indonesia's coal reserves in 2010 amounted to 21000 million

tonnes. When viewed from the coal reserves, estimated coal reserves in Indonesia

will expire 71 years from now. The estimates based on assumptions about the

level of production 391 million tonnes per year and not new reserves are found

(Komaidi, 2013). In addition, the development of coal production over the past

few years continues to show increased rapidly, with an average production

increase of 15.68% per year. It makes the process of coal gasification becomes

very feasible conducted as a raw material of ethylene production.

1.2. Review Literature

1.1.1. EthyleneEthylene is lightest hydrocarbon olefin (double chain) with a molecular

weight of 16, colorless, flammable, and slightly fragrance. The properties of

ethylene is determined based on double bond, the main reaction is an addition

reaction produces saturated hydrocarbons and their derivatives or polymers (Kirk

& Othmer, 1977). Now, almost all ethylene made from natural gas, ethane,

propane, and other paraffin and heavy fractions of crude oil, naphtha, kerosene,

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and gas oil. A small amount of ethylene gas obtained from refinery output

(catalytic cracking). Some ways of making ethylene according to Mc. Ketta

(1984) are:

a. Hydrocarbon Pyrolysis

This technique is most widely used in the manufacture of ethylene. The

reaction equation of hydrocarbon pyrolysis is:

C7H16 ½ C5H12 + 1/3 C4H8 + 1/3 C3H8 + 1/3 C4H10 + 1/3 C3H

Large-scale production is done by hydrocarbon raw material pyrolysis and

added with the ratio in a steam heater and pyrolysis continued by separation of the

resulting mixture of gas passing through a complex operating system. This

process is produces a complex mixture of hydrocarbon products and will be more

complex again as more with the weighs of the hydrocarbon molecules in the

pyrolysis.

High yield selectivity toward the desired olefins and olefins (ethylene,

propylene, butadiene), methane results, and minimum coking can be achieved by

operating the heater at high temperature pyrolysis (750-900˚C), short residence

time, and low partial pressures. The addition of steam serves to reduce the

hydrocarbon partial pressure and the amount of carbon deposited in the tube wall.

The weight ratio of steam to hydrocarbon varies from 0.3 to 1.0 for ethane to gas

oil. Changes in the carbon chain paraffinic and naftenik to olefins.

b. Ethanol Dehydration

Manufacture of ethylene from ethanol dehydration followed the following

equation:

C2H5OH C2H4 + H2O

The reaction occurs with activated alumina catalyst and phosphoric acid.

Ether formation occurs at a temperature of 230˚C, while at temperatures 300-

400˚C ethylene obtained with a minimum content of ether. Ethylene yield can

reach 94-99% from the theoretical values depending on the process used. Further

purification is used to separate acetaldehyde, acids, other hydrocarbons, CO2, and

water. The process is developed on a small scale in Europe, America, and

Australia in the 60's, before the development of the ethylene plant produces

ethylene cheaper, that of hydrocarbons.


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c. Propylene Dispropotionation

In this process, which is relatively cheap propylene converted into ethylene

and butylene higher price with the help of tungsten oxide-silica catalyst.

2C2H6 C2H4 + C4H8

d. Ethylene From Coal

An indirect manner and process of considering alternative oil and natural

gas supplies are running low while coal still more. This method involves three

processes, namely:

Production of synthesis gas from coal gasification process (process Lurgi,

Koppers-Totzek, Winkler).

Synthesis gas is converted to hydrocarbons by the Fischer-Tropsch process.

Ethylene is made by pyrolysis of hydrocarbons or hydrated ethanol is obtained.

This process runs in South Africa that do not contain petroleum but rich in

coal. From all the above, the most widely used is a hydrocarbon pyrolysis process,

which consists of three stages, is :

Synthesis.

Recovery.

Purification.

The process is carried out at high temperature between 1500 - 2000˚F.

Techniques used in the production of ethylene by means of pyrolysis of

hydrocarbons in general according to Kirk & Othmer (1977), among others:

Fired tubular heater.

Regenerative stove.

Moving bed / refractory pebbles.

Generation of heat by combustion in air or oxygen use.

The electric arc.

Contact with tin.

Both the first-mentioned process, ie the process by using tubular and

regenerative heater fired stove has been run commercially. Of the six above

process, this time almost all the ethylene plant using tubular fired heater in its

production. Recovery and purification of ethylene is a complex process, given the


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results of pyrolysis is also very complex. The process is run commercially for this

purpose are :

Fractionation at low temperatures and high pressures.

Adsorbtion and fractionation at low temperatures.

Fractionation at low temperature and low pressure.

Extraction solvent

For the record fractionation process at low temperatures and high pressures

are used in 75% ethylene plant operating in the United States. In the design of this

ethylene plant licensing process there are 3 most dominant and widely used,

namely Kellogg, Lammus, and Linde. Among the licensor third, Kellogg process

with the front end demethanizer scheme that was developed around 1960 had an

estimated 40% of all existing plant in the world, due to the high thermodynamic

efficiency.

1.1.2. CoalCoal is a fossil fuel that its formation requires certain conditions and in a

long time. Coal is the remains of plants that change shape, which originally

accumulated dirawa and peatlands. Hoarding silt and other sediments, together

with a shift of the earth's crust (known as tectonic shifts) swamps and peat buried

to a depth that is often very deep. With such hoarding, plant material is exposed to

high temperature and pressure and causes the plants undergo a process of physical

and chemical change and transform the plant into peat and then coal, a process

known acoalification (Bell, 2011).

Quality of each coal deposit is determined by varying the temperature and

pressure as well as the length of time the formation of the so-called "organic

maturity". Initially the peat is converted into lignite (bararmuda stone) or brown

coal (brown coal). Influenced by temperature and pressure constant, a young coal

which changes gradually add organic maturity and young converts coal into a sub-

bituminous coal. Chemical and physical changes continued until coal becomes

harder and the color is black and form a bitumen or antrasi. Under the right

conditions, an increase in the organic maturity continue, finally forming

anthracite.


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1.1.3. Types of Coal and Its ApplicationFigure below shows the classification of coal. It can be seen that the coal

generally is divided into two main groups, namely low quality coal or soft coal

and hard coal. This classification is based on the carbon content and the humidity,

the more the carbon content of the coal, the better. While it is inversely

proportional to the moisture of coal, the coal, the coal moisture smaller the better.

Figure 1.1. Coal and Utilization(Source: Research and Development Center for Mineral and Coal Technology, 200B, 2006)

Low quality coal there are two types, namely lignite and sub-bituminous.

Coal is usually more gentle with fragile materials such as soil and gloomy color.

This coal has a relatively high humidity levels and low carbon content so that the

content of energy or low calorific value. This type of coal utilization for power

generation typically, the source of energy for industry, and cement production.

While coal with higher quality or hard coal are generally more powerful and often

brilliant black like glass, the higher carbon content and lower moisture levels

make this coal has a higher heating value. Anthracite is the most excellent quality,

having carbon content and higher energy and lower humidity levels. Utilization of

coal types are also more widely, in addition to power plants and cement industry,

coal is also used in the manufacture of iron and steel industry.


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1.1.4. Coal GasificationCoal gasification is the process of converting the carbon in the coal to gas

(syngas) using gasification medium (gasification agent). Substance used as a

medium is air or oxygen, water or steam, and carbon dioxide. The gas produced

diverse as carbon monoxide and hydrogen which is the main product, carbon

dioxide, steam, methane and even gases such as NO% and SO%, although in

small amounts. Stages in the gasification process can be seen in the picture below

along with an explanation.

Figure 1.2. Coal Gasification Process (Source: www.esptk.fti.itb.ac.id, 2010)

a. Drying (Evaporation of Moisture)

Phase Drying (Evaporation of Moisture) Is the stage where the evaporation

of water content in the coal. This process depends on the type of coal,

subbituminous coal to the type of humidity can reach 35% by weight, while the

smaller bituminous about 5% by weight. This process can be avoided in the

gasifier by means treatment first coal input (dehydration).

b. Pyrolysis

Is a chemical process of decomposition by means of heating at minimal or

no oxygen conditions. Coal will experience cracking on this process at

temperatures of about 250oC to above. The result of this process is charcoal, water

vapor, tar vapors, and gases. The purpose of this stage is to facilitate the formation

of syngas at the gasifier reaction. In the pyrolysis process resulting CO, CO2, CH4,

H2O, H2, and tar.


http://esptk.fti.itb.ac.id/

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c. Formation of Volatile Substances (Combustion of Volatile Matter)

Substance reacts with oxygen pyrolysis results to generate the heat required

by the next process, gasification by steam. The process of oxidation (burning) is a

homogeneous reaction is exothermic and provide heat or energy for

heterogeneous reaction process. The reaction below is a common homogeneous

reaction.

C ( s )+O2 ( g ) →CO2 ( g )∆ H rxn° =−393.98 kJ /gmole

2 CO ( g )+O2 (g )→ 2CO2 (g ) ∆ H rxn° =−566.65 kJ / gmole

2 H 2 ( g )+O2(g)→2 H 2O ( g ) ∆ H rxn° =484.23 kJ /gmole

d. Gasification (Heterogenous Reaction)

At this stage, the bias char reacts with water vapor and carbon dioxide to

produce hydrogen and carbon monoxide gases as the main component. At this

stage is the stage that occurs digasifier or gasification processes. This process is

dependent on the selection of the gasifier itself, where each gasifier it has

advantages and disadvantages. The following reactions occur in this process.

C ( s )+H 2O (g )→ CO ( g )+H 2 ( g ) ∆ H rxn° =+131.46 kJ / gmole

C ( s )+CO2 (g )→ 2CO ( g ) ∆ H rxn° =+172.67 kJ / gmole

C ( s )+2 H 2 (g )→ CH 4 (g ) ∆ H rxn° =−74.94 kJ /gmole

At this stage also formed methane in certain circumstances. To adjust the

ratio of H / C to fit the needs, usually arranged so that the case of water-gas shift

reaction reaction that aims to convert CO to H2.

CO ( g )+H 2O (g )↔ CO2 ( g )+H 2 (g ) ∆ H rxn° =−41.21 kJ / gmole

e. Slagging

Inside there is also a coal of ash content (ash) in which the amount depends

on the type of coal and the coal geography. This ash is separated by slagging

process.

1.2. Market AnalysisIn planning the construction of a plant, in addition to the availability of raw

materials are cheap and easily should also be noted that the development of the

market of goods to be produced in this case is a short chain olefins ethylene.


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Therefore, it takes a market analysis which includes the development of

production, export, import and consumption.

Based on the data table below, the development of ethylene production in

Indonesia has decreased the period 2005 – 2009. In contrast, domestic

consumption has increased an average of 7.8% per year from 2005 to 2009

reached 1,118,000 tons capacity so that the necessary import as much as 660,000

tons in 2009 to meet the needs primarily of ethylene as a feedstock for the

production of ethylene derivatives such as polyethylene, vinyl acetate, ethylene

oxide, ethyl benzene, and so forth. With conditions that increased imports for

domestic needs, is no longer possible to export abroad.

Table 1.1. Data Capacity Production, Consumption, Exports, Imports Ethylene Year Period 2005-

2009

Capacity x 1000 ton 2005 2006 2007 2008 2009

Production 510 460 540 488 455Konsumsi 847 754 801 931 1118Ekspor 0 0 0 0 0Impor 337 294 261 444 660(Source: Ministry of Industry, 2011)

Based on data from each of the five years from 2005 to 2009 made the

prediction capacity of domestic production and consumption for the period 2010

until 2025 and depicted in Table 1.2. and Figure 1.3. As for the import of data

from the years 2010 to 2025 is the difference of the capacity of production and

consumption each year from predicted results. From the results obtained estimates

of data that needs to ethylene will continue to increase until it reaches a capacity

of 11,171,000 tonnes, while production in the year only 2,263,000 tonnes that will

be required for the import of 8,908,000 tonnes outside.

Table 1.2. The Prediction Capacity Production, Consumption, and Imports Ethylene Year Period

2010-2025


10

Year

Consumption

Production

Import

(tonne/year)

(tonne/year)

(tonne/year)

2010 1109000 468000 6410002011 1296000 460000 8360002012 1459000 430000 10290002013 1600000 427000 11730002014 1754000 421000 13330002015 1999000 405000 15940002016 2374000 481000 18930002017 2820000 571000 22490002018 3350000 679000 26710002019 3978000 806000 31720002020 4725000 957000 37680002021 5613000 1137000 44760002022 6667000 1351000 53160002023 7918000 1604000 63140002024 9404000 1905000 74990002025 11171000 2263000 8908000

(Source: Author’s Personal Data)

20102012

20142016

20182020

20222024

0

2000000

4000000

6000000

8000000

10000000

12000000

ConsumptionProductionImport

Years

Capa

city,

tonn

e/ye

ar

Figure 1.3. The Prediction Capacity Production, Consumption, Imports(Source: Author’s Personal Data)


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Of the real data and the predictions that have been discussed previously

showed an imbalance between production capacity and domestic consumption.

Domestic production capacity is still very small compared to consumption

necessitating an increase in ethylene production capacity in the country. In 2012,

producers of ethylene producer, PT. Chandra Asri produce ethylene with a

capacity of 600,000 tons and PT. Pertamina with a capacity to produce 405,000

tons of propylene. With two of the manufacturers producing olefins, it is an

opportunity for the establishment of an ethylene plant can help cover the

imbalance of production and domestic consumption will be a short chain olefins.

1.3. Capacity AnalysisThe production capacity of a plant can be determined by

market analysis that have been made. Market analysis indicates

any opportunities that can be taken by the new manufacturers to

invest in ethylene industry. Opportunities can be maximized

among the competition with other manufacturers, both the

existing manufacturers and new manufacturers. This opportunity

is greater due to higher economic growth in Indonesia which

always above 5 percent in the last ten years. Stable economic

growth for years will be followed by growth in the industrial

sector as a result of the increasing ability of consumers to buy

the commodities.

Capacity analysis is needed to determine the production

capacity according to manufacturers's ability. Although there is

an opportunity to benefit from the gap between demand and

supply, there should be anticipation of the possibility of

competition between manufacturers of ethylene. Rivalry can be a

competition with manufacturers which had already producing

and new competitors that could emerge later on.

Based on the analysis of the ethylene market, ethylene

demand in Indonesia has grown on year to year. Although


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ethylene demand growth followed by growth in ethylene

production, but the amounts are relatively insignificant. This is

illustrated by the persistence of the difference between the

demand and supply of ethylene in Indonesia. Market analysis

concludes that there is a difference or deficit between supply and

demand for ethylene in Indonesia amounted to 2,671,000 tons

per year in 2018.

That figure could not necessarily be used as a new

ethylene capacity of a plant to be built. Since the existing

manufacturers are still very likely to expand their production

capacity and the possibility of new competitors in the coming

years. The existing existing manufacturers could produce

ethylene ranged from 250,000 tons to 550,000 tons.

Based on the figure of deficit of ethylene supply and the

existing manufacturers production capacity, our factory will

produce 400,000 tons of ethylene per year. That figure is

equivalent to the production of 1212 tons of ethylene per day.

Our production capacity is equal to 30 percent of the gap

between supply and demmand of ethylene in Indonesia. The rest,

70 percent, is our anticipation for the probability of rivalry.

We are very confident that the production capacity of

400,000 tons of ethylene per year will be able to be absorbed by

the market. Considering the ethylene consumption which is

continously increasing every year, the gap of 896,000 tonnes will

increase following the growth of demand for ethylene. Production

capacity of 400,000 tons is also the most ideal figure to get some

profits. Due to the growing production capacity, it is better to get

a category of economically feasible before build the plant and

consider the possibility of competition with competitors.


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1.4. Raw Material AnalysisCoal is a fossil fuel that it’s formation requires certain

conditions and lasts within a long time. Coal is the remains of

plants that deformed and originally accumulated in swamps and

peat bogs. Stockpiling of silt and other sediments, along with a

shift of the earth's crust (known as tectonic shifts) burried

swamps and peat to a depth which is often very deep. With such

piling, plant material is exposed to high temperature and

pressure and causes the plants undergo a process of physical

and chemical change and transform the plant into peat and then

coal, now it is called the coalification process.

One of the indicators of the quality of coal is the content of

carbon, these are the carbon composition for the five types of

coal:

Anthracite coal is the highest grade, with glittering black color

(luster) metallic, containing between 86% - 98% of the

elements of carbon (C) with a water content of less than 8%.

Bituminous containing 68-86% of carbon element (C) and

water content of 8-10% by weight. This Grade of coal is the

most mined in Australia.

Sub-bituminous contains less carbon and more water, and

thus it's a source of heat which less efficient compared to

bituminous.

Lignite or brown coal is very soft coal which contains 35-75%

water by weight.

The quality of coal is also determined by the temperature

and pressure as well as the length of time for the formation of

the so-called organic maturity. Initially the peat is converted into

lignite (lignite) or brown coal. This is the type of coal with low

organic maturity. Influenced by temperature and pressure

constant, a young coal which changes gradually add organic


14

maturity and young coal converts into a sub-bituminous coal.

Chemical and physical changes continue to occur until these

coals became harder and the color is black and form a bitumen

or antrasi. Under the right conditions, an increase in the organic

maturity can continue, finally forming anthracite. Figure 1.4

shows coal development process.

Figure 1.4. Coal development process(Source : Habiburrohman, 2012)

There are two types of low-quality coal, the lignite and sub-

bituminous. The coal is usually more gentle with fragile materials

such as soil and gloomy colored. The coal is young, has a fairly

high moisture levels and low carbon content, thus the energy

content or heating value is low, the use of the average of coal is

to generate electricity, industrial and energy sources for the

production of cement.


15

Figure 1.5. Coal Quality Based on Coal Type(Source : Habiburrohman, 2012)

Coal with higher quality or hard coal are generally more

powerful and often brilliantly black like glass, the higher content

of carbon and lower moisture levels make this coal has a higher

heating value. Anthracite is the most good quality of carbon

content and thus it has higher energy and lower humidity levels.

The use of coal is greater, in addition to power generation and

cement industries, coal is also used in the manufacture of iron

and steel industry.

By 2010, the average selling price of coal at the mines that

produce each of the four main types of coal are :

Table 1.2. Average Coal Price According to The Type

Coal TypePrice per Ton

($)

Lignit 18.76

Sub-bituminuos 14.11

Bituminous 60.88

Anthracite 59.51

(Source : www.esdm.go.id)

The selection of coal based on several parameters such as

the composition of carbon, the price and hardness level. Followed

by an assessment of each type of coal to get the type of coal that


16

will be used. Below is a table of assessment for the selection of

raw materials :

Table 1.3. Raw Material Selection

Raw

Material

The

Composit

ion of

Carbon

PriceCoal’s

Hardness

Total

Score

Lignit -1 +1 -1 -1

Sub-

bituminous0 +1 -1 0

Bituminous +1 0 +1 +2

Anthracite +1 0 -1 0

(Source : Author’s Personal Data and gather from any sources)

Bituminous coal has several advantages when compared

with other types of coal. Bituminous coal has high carbon

content. Although the carbon content in the anthracite coal is

higher, but it also harder. Consequently size reduction process is

more difficult when using anthracite coal. It means more time

and bigger cost to reduce the coal size. While the sub-bituminous

coal and lignite contains less even very low carbon. The water

content in the lignite and sub-bituminous coal also very high.

Coal with a low carbon content and high water content will

produce ethylene in a very low yield. So that ethylene production

using lignite and sub-bituminous need more coal when compared

to ethylene production with bituminous coal. Although the price

of bituminous coal is more expensive than the two types of coal,

but it will actually save storage space and transport costs since

less coal is needed.

Considering the carbon composition, price, and hardness

then bituminous coals are selected. While coal suppliers selected

from PT Bukit Asam and PT Adaro which is located in southern

Sumatra.


17

Bituminous coals have composing components with the

following composition :

Table 1.4. The Composition of Bituminous Coal

Component % Weight

Moisture 1.0 – 10.0

Ash 4.0 – 20.0

Sulfur 0.5 – 2.2

Table 1.4. The Composition of Bituminous Coal (Cont’d)

Component % Weight

Fixed carbon 50.0 – 72.0

Volatile matter and

another elemental and

compounds (H2, N2, and

Cl2)

17.0 – 37.0

(Source : Habiburrohman, 2012)

Here is physical properties for bituminous coal :

Table 1.5. Bituminous Coal Physical Properties

Physical Properties Data

Igniton temperature 260o – 365oF

Melting point 750oF

Average specific

gravity1.43

% volatility by volume Negligible

Vapour density (air = 1) N/A

Solubility in water Non-soluble

(Source : MSDS of Bituminous Coal)

Since the gasification unit need oxygen feed, we have to get

oxygen supply. We preferred to have a partnership with a third

party. Thus, it is not necessary for us to build the unit. They will

build their own air separation unit as oxygen supply in our site. It

is more efficient than we build the unit ourself.


18

Oxygen can be produced from air. The air will be processed

by fractination to seperate the component such as nitrogen,

oxygen even helium. The fractination components of air is done

by cryogenic temperature. The air will be condensed or

liquifaction and then will be separate with cryogenic distillation to

seperate all component with high purity. The process can use

zeolite as membran filter to improve the purity of gas. Oxygen

with high purity, 90% to 93% can be produced from this process.

Here is physical properties for oxygen :

Table 1.6. Oxygen Physical Properties

Physical Properties Data

Molecular weight 32 g/mole

Moleculr formula O2

Boiling/condensation

point

-183oC (-297.4oF)

Melting/freezing point -218.4oC (-361.1oF)

Critical temperature -118.6oC (-181.5oF)

Vapor density 1.105 (air = 1)

Specific volume 12.0482

Gas density 0.083

(Source : MSDS of O2)

The supply of oxygen can be fulfilled by some companies such as PT. Air

Liquid and PT. Air Product. We preferred to PT. Air Product Indonesia,

since they has technology which called oxygen production on-site

and mostly produce gaseous oxygen. The technology which used

by PT. Air Product Indonesia is PRISM. PAT. Air Product Indonesia

will build a oxygen supply facilites in our plant which will be

managed by PT. Air Product Indonesia. The oxygen supply

facilites will be integrated with our process plants. This


19

technology has low investment and low operating cost. This

technology is economical alternative to supply oxygen for our

plants by cryogenic treatment. The range of purity from this

technology is 90%-93% and the capacity is 200-10.000 Nm3/hour.

1.5. Plant Location AnalysisThe selection of appropriate location of the plant will

adversely impact the operating and capital costs which is should

be inculcated by investors to build a plant. It will directly affect

the economic value and feasibility of the plant to be built. To

choose the right location, there are several factors of

considerations that must be considered. Because of these factors

will affect the continuity of the production of the plant in the

future.

We will build a plant in Cilegon, Banten. Here is the

candidate location for this plant in Cilegon :

Figure 1.6. Plant Location(Source : www.maps.google.com)

The selection of the location of the plant in Cilegon is

determined by various factors of considerations, such as:


20

a. Raw Material Availability

Our factory’s main raw material is the coal. So the proximity of this plant

with the areas rich in coal reserves that have been explored is crucial. Here is a

map of the distribution of coal reserves in Indonesia:

Figure 1.7. Indonesia’s Coal Reserve(Source : www.esdm.go.id)

From the map above, the southern part of Sumatra and

Borneo is the the area with the largest coal reserves and

production in Indonesia. Although Cilegon not really close to

Borneo and Sumatra, but closer to the geographical location of

southern Sumatra. The coal will be used as raw material in this

plant comes from PT. Bukit Asam and PT. Adaro, these

companies are in the province of South Sumatra.

b. Target Market Location

Petrochemical company is widely built in Java, especially

around Cilegon and Greater Jakarta. So the site selection in

Cilegon make this plant location closer to the consumer and not

really far from coal mine in South Sumatra. Here is an illustration

of strategic Cilegon if viewed in terms of the location of the

target market and the location of the coal mine :


http://www.esdm.go.id/

21

Figure 1.8. Target Market, Coal Mining and Plant Location(Source : www.maps.google.com)

c. Infrastructure and Transportation

Cilegon is now connected with the Jakarta-Merak Toll Road.

So access to and from the Greater Jakarta is very easy. In

addition, the location of Cilegon which is on the northwest coast

of Java island is an area that lives within range of industrial port

of IPC II. So it will be very easy transporting coal by sea.

d. Workers Availability

As an industrial city, Cilegon has a reserve of human

resources reliable in quality and quantity. By 2014, there are

approximately 398,304 inhabitants in Cilegon with a workforce of

60,000 people. In addition, the availability of labor from the area

around and outside Cilegon is also relatively large. It can be seen

from the number of workers from several areas in Java such as

Greater Jakarta, West Java, Central Java and East Java.

e. Government Policy and Sosio-Geographical Circumstance

Cilegon social circumstances today are very heterogeneous.

Nevertheless hardly ever there is a riot on a large scale in

Cilegon within the last few decades. This condition makes


http://www.maps.google.com/

22

Cilegon has been chosen by many companies to set up their

factories.

By 2014, the MSE is Rp2.443.000 in Cilegon. That number

are considered rational for the petrochemical industry because of

the labor in this sector is mostly skilled and trained workers.


CHAPTER IIPROCESS SELECTION

2.1. General ProcessThe goal of this plant is convert coal become ethylene. Figure 2.1. shows

black box for this plant.

Syngas EthylenePropyleneEthanePropaneButhenePentheneHCOOHH2OMethaneMethanol

Figure 2.1. Black Box of This Plant(Source : Author’s Personal Data)

Input for this plant is syngas. The input will be processed in black box and

the result is ethylene, propylene, ethane, propane, buthene, penthene, HCOOH,

H2O, methane, and methanol. This process will be started in gasification process.

In this gasificication process, coal will be converting become syngas first. This

process will be continued to synthesis of ethylene and the result is ethylene.

2.2. Alternative ProcessProcess production ethylene from coal has many kind of altrnative process.

Study literature to find available process is needed which can be used to produce

ethylene from coal. First, coal will be gasified and produce synthesis gas. This

product will be fed to the next process to synthesis olefin such as ethylene.


BLACK BOX

24

2.2.1. Gasification TechnologyCoal gasification is a series of reaction steps that convert coal containing C,

H, and O, as well as impurities such as S and N, into synthesis gas and other

forms of hydrocarbons. Synthesis gas is mixture between from hydrocarbon gas

and carbon monoxide that convert’s result from coal.

In total gasification, coal reacts with solvent, consist of air, oxygen, steam,

carbon dioxide or mixture of the gas. All organic materials from coal will be

changed become gas product so residue in the reactor is dust. The main reaction

that happen in the gasification process by using air is called coal partial

combustion. Gas product that resulted is also called producer gas or low calorie

gas because it contains carbon monoxide and nitrogen also the heat value of this

gas approximately 1780 calorie/litre.

This conversion is generally accomplished by introducing a gasifying agent

(air, oxygen, and or steam) into a reactor vessel containing coal feedstock where

the temperature, pressure, and flow pattern (fixed bed, fluidized bed, or entrained

bed) are controlled. The proportions of the resultant product gases (CO, CO2, CH4,

H2, H2O, N2, H2S, SO2, etc.) depend on the type of coal and its composition and

the gasifying agent (or gasifying medium).

The composition of gas that resulted from gasification process depends on

feedstock (type of coal are used), operation temperature, and influence of steam.

At the high temperature, gas that resulted contains rich carbon dioxide, whereas at

the low temperature, gas that resulted contains rich carbon monoxide. Moisture in

coal or steam that added into reactor could increase hydrogen and carbon

monoxide proportion in the gas product and heat value of gas. In case excessive

water or steam that added, process temperature would decrease resulted rich

carbon dioxide effected low heat value.

The next step after become syngas is acid gas removal. In high temperature

processes, all sulfur components in the feed are converted to H2S or COS, which

are undesirable components and require acid gas removal technology to be

treated.

Actually, water gas shift reaction is not classified as one of the principal

gasification reaction. It cannot include in the analysis of chemical reaction that


25

involve syngas. Among all reactions involving syngas, so equilibrium constant is

least strongly dependant on the temperature.

Syngas product from a gasifier contains a variety of gaseous species such as

carbon monoxide, hydrogen, carbon dioxide, methane, and water (steam). Water

gas sfift reaction is one of the major reactions in the steam gasification process,

both water and carbon monoxide are presents.

2.2.4. Synthesis Ethylenea. Method 1 –Dehydration of Methanol

Methanol to olefin production starts from syngas conversion into methanol.

Synthesis gas (syngas) is a mixture that contains hydrogen, carbon monoxide and

carbon dioxide as principal components and methane and steam as secondary

components. Methanol synthesis from syngas commonly conducted over

CuO/ZnO/Al2O3 catalyst. In this type of catalyst, Al2O3 plays role as a support.

The reactions involved in methanol synthesis from syngas are:

CO+ H 2⟷CH 3OH

CO2+3H 2 ↔CH 3 OH+H 2

CO+3H 2O ↔CH 3OH +H 2 O

There are several side reactions that can occur beside the reactions

mentioned above. One of the main byproducts that produced in methanol

synthesis is dimethyl ether. The reaction of dimethyl ether synthesis can be

written below :

2 CH3 OH ⇔ CH 3O CH3+H 2 O

The typical methanol synthesis technology involves several steps, namely

syngas compression, catalytic synthesis, crude methanol distillation and recycle &

recovery. There are several well-known methanol synthesis technology, such as

Conventional ICI’s 100-atm Methanol Process, Haldor Topsoe A/S Low-Pressure

Methanol Synthesis Process, Kvaerner Methanol Synthesis Process, Krupp

Uhde’s Methanol Synthesis Technology, Lurgi Oel-Gas-Cheme GmbH Process,

Syntetix LPM Process and Liquid Phase Methanol Process.

The Conventional ICI’s 100-atm Methanol Process was first developed in

1972 when Cu/ZnO/Al2O3 catalysts system was announced. This process

originally consists of two section, reforming and synthesis. Haldor Topsoe A/S


26

Low-Pressure Methanol Synthesis Process is designed to produce methanol from

natural or associated gas feedstocks, including a two-step reforming process to

generate syngas mixture for methanol synthesis. Kvaerner Methanol Synthesis

Process is similar to Haldor-Topsoe process, the difference is that in this process

carbon dioxide can be used as a supplementary feedstock to adjust the syngas

ratio. Krupp Uhde’s Methanol Synthesis Technology has flexibility of feedstock

choice, including natural gas, liquefied petroleum gas or heavy naphta. Lurgi Oel-

Gas-Cheme GmbH Process is meant to produce methanol in a single-train plant

with feedstocks from natural gas or ail-associated gas. Meanwhile, Syntetix LPM

Process is an impoved version of ICI’s technology. This process is designed to

handle natural gas feedstock, but also capable of handling other hysdrocarbon

feedstock like naphta, coal and other petrochemical offgas stream. The last

commercial technology for methanol production is Liquid Phase Methanol

Process. In this process, the reaction carries out in a slurry reactor using

Cu/ZnO/Al2O3 catalyst.

Methanol is produced from the process before can be used to synthesis

olefin. Product from this process is not only ethylene, but also propylene.

Methanol to olefin (MTO) enables low costs of production ethylene (and

propylene) in regions that do not have large resrves of ethane. MTO is used in

some country to fill the gap between ethylene demand and supply.

The conversion of methanol to olefins and other hydrocarbons products has

been widely studied. This method is developed and licensed by UOP. This process

use SAPO-34 catalyst, silicoaluminumphosphates. SAPO-34 has unique pore

size, geometry and acidity created a more selective route for methanol conversion

to ethylene and propylene with reduced heavy by product.

MTO process utilises a fluidised reactor and regenerator system to convert

methanol to olefins using a proprietary, SAPO-34 Catalyst. MTO process can be

operated on crude, or undistilled methanol, as well as with pure methanol. The

choice of feedstock quality generally depends on project spesific situations

because there can be advantages in either case. Figure 2.2. ilustrates a simple flow

diagram for the MTO process by UOP.


27

The methanol feed is preheated and then introduced into the reactor. The

conversion of methanol to olefins requires a selective catalyst that operates at

moderate to high temperatures. The reaction is exothermic so heat can be

recovered from the reaction. Carbon aor coke accumulates on the catalyst and

requires removal to maintain catalyst activity. The coke should be removed by

combustion with air in a catalyst regenerator system. A fluidised bed reactor and

regenerator system is ideally suited for the MTO process because it allows for

heat removal and continuous catalyst regeneration. The rractor operates in the

vapor phase at temperature 650-1000oF and pressure between 15 and 45 psig. A

slipstream of catalyst is circulated to the fluidised bed regenerator to maintain

high activity.

The reactor effluent is cooled and quenched to seperate water from the

product gas is compressed and then unconverted ixygenates are recovered and

returned to the reactor. The reactor provides very high conversion so there is no

need for a large recycle stream. After the oxygenate recovert section, the effluent

is further processed in te fractionation and purification section to remove

contaminants and separate the key products from the by product components.

Ethylene and propylene are produced as polymer grade products and sent to

storage. The havier component can be sent to OCP reactor where it is selectively

converted to light olefins.

Figure 2.2. Simple Flow Diagram for the MTO process by UOP(Source : A, Grefory, et.al. Hydrocarbon Engineering)


28

b. Fermentation and Dehydration of Bio-Ethanol

Ethanol to olefin is a similar process to methanol to olefin. Syngas produced

in the gasification process can be converted into ethanol using metal catalyst of

microbial catalyst. Currently, syngas conversion to ethanol using metal catalyst is

used in industries. The types of metal-based catalyst for producing ethanol from

syngas are Rh-based catalyst. However, this technology is limited in selectivity of

conversion and some catalysts used are sensitive to poisoning. These catalysts are

also required severe condition with complex composition in producing ethanol.

To overcome this problem, microbial agent is developed to convert syngas

to ethanol. Several bacteria are able to convert syngas into ethanol as part of their

metabolism. The overall stoichiometric reaction of ethanol production from

syngas is shown below :

6 CO2+3 H 2 O→ CH 3CH 2 OH+4CO2 (1)

2 CO2+6 H 2 →CH 3 CH2 OH+3 CO21 (2)

6 CO+6 H 2→ 2 CH3 CH 2OH+2CO2 (3)

Microbials that are able to convert syngas into ethanol belong to acetogen

class. Acetogens are microorganisms that can use gases like CO2, H2, and CO as

well as other substrate like sugar to produce ethanol, acetate and cell carbon via

acetyl-CoA pathway.

There are several kinds of reactors that can be used for processing syngas

fermentation to ethanol. Trickle-bed reactor (TBR) that consists of vertical tubular

reactor can be used for fermentation by attaching the microorganisms into packed

solid material. Other reactor that can be used for fermentation is continuous

stirred-tank reactors (CSTR). In CSTR, a continuous flow of gas bubbling through

the liquid which consists of a dilute solution of essential nutritions needed by the

microorganisms to grow and survive. Beside TBR and CSTR, packed-bed reactor

(immobilized-cell reactors) also can be used for syngas fermentation. This reactor

usually consists of columns packed with microorganisms that are immobilized

into biocatalyst reactor.

The production of ethanol from syngas basis using microbial agent is

currently not available on a commercial basis. This technology has disadvantages

in reaction time of water-gas shift reaction due to the slow cell growth of the


29

microorganisms. The fermentation is also often limited by low productivity and

the rate of syngas that can be transferred into the liquid.

Meanwhile, the ethanol dehydration to produce ethylene usually consists of

three sections, which are reaction and quenching section, the compression, caustic

washing and drying section, and the purification section. The reaction occurs in

four stages of adiabatic, fixed-bed reactor. The output from the last reactor is then

quenched to reduce its water content. After quenched, the output is compressed

and sent to as caustic washing column to reduce CO2 content. This stream is then

dried in a drying system that consists of two molecular sieve beds. The stream is

separated into vapor and liquid components before being fed into two different

ethylene columns to remove C3+ residues from the dehydration reaction. The

overhead product that consists of ethylene is then undergoes CO removal in a

stripper column.

c. Fischer Tropsch

Fischer-Tropsch process can also be used for converting syngas produced in

coal gasification process into olefins. Fischer-Tropsch is a reaction that converts

syngas that consists of carbon monoxide and hydrogen into liquid hydrocarbon.

Syngas produced in gasification process is purified first from impurities like

hydrogen sulfide, carbon dioxide, tars and slag, before entering the Fischer-

Tropsch reactor. There are three main reactions and two side reactions in Fischer-

Tropsch reactor :

Main reaction

Alkanes nCO+(2 n+1 ) H 2→ Cn H 2 n+2+n H 2 O

Alkenes nCO+2 n H 2→ Cn H 2n+n H 2O

Water-gas shift CO+ H 2O⇆CO2+ H 2

Side reaction

Alcohol nCO+2 n H 2→ H ¿¿

Boudouard reaction 2 CO→ C+C O2

The product from the reactor varies according to the reaction temperature,

catalyst, pressure and syngas composition (H2/CO ratio). The products then

undergo separation to separate the fractions produced from the reactions. The


30

hydrocarbon fraction produced in the reactor includes ethylene, propylene, LPG,

naphtha, and longer hydrocarbon chain fraction. The product that is used for

olefin production is naphtha fraction, so other fraction than ethylene, propylene

and naphtha need to be reformed into naphtha first.

The formed naphtha is then being cracked to produce olefins. The overall

cracking process of naphtha can be seen in the figure below :

Figure 2.3. General Block Diagram Flow of Fischer Tropsch(Source : www.fisher.com)

The first step of the olefin production is cracking naphtha into ethylene and

other various products in a furnace. This process is called pyrolysis, which can be

defined as thermal cracking of petroleum hydrocarbon with steam. The cracked

gas is then quenched to preserve the current gas composition and prevent

undesirable side reactions. After being cooled, the cracked gas is then compressed

in a turbine driven centrifugal compressor with typically four or five stages and

interstage cooling. Prior entering the compressor and after each interstage cooling,

the gas is dried to prevent the formation of hydrates and ice. Between 3rd and 4th or

4th and 5th stages, acid gas removal system is also placed to prevent formation of

hydrates and ice in the next step. The compressed gas is the fractionated in

distillation column to produce different products and fractions at specified

qualities.

2.3. Process SelectionIn this sub bab, we would select the gasification technology that we use.

This selection consider many variables (based on the literature), such as type of

coal that used (coal rank) as raw material, efficiency from energy side and utility


http://www.fisher.com/

31

(total of gasification media), also capital cost from technology that we used,

conversion of carbon, and purity of syngas as output from gasifier. In this sub bab,

we also select the acid gas removal technology and continous process from syngas

resulting ethylene.

2.3.1. Gasification Technologya. Selection of Gasification Process based on Reaction

Steam Gasification

The steam gasification reaction is endothermic, so this process requiring

heat input for the reaction to proceed in its forward direction. Usually, an excess

amount of steam is also needed to promote the reaction. However, excess steam

used in this reaction hurts the thermal efficiency of the process. Therefore, this

reaction is typically combined with other gasification reactions in practical

applications. The H2-to-CO ratio of the product syngas depends on the synthesis

chemistry as well as process engineering. Reaction of steam gasification is shown

below.

C(s )+H 2O(g)↔CO(g)+H2(g )∆ H298o =131.3kJ /mol(2.1)

In the case of catalytic steam gasification of coal, carbon deposition reaction

may affect the catalyst’s life by fouling the catalyst active sites. This carbon

deposition reaction is more likely to take place whenever the steam concentration

is lacking.

Carbon Dioxide Gasification

The reaction of coal with CO2 is endothermic, similar to the steam

gasification reaction. Reaction of carbon dioxide gasification is shown below.

C(s )+2CO2(g )↔ 2CO(g)∆ H 298o =172.5 kJ /mol (2.2)

The reverse reaction is a carbon deposition reaction that is a major culprit of

carbon fouling on many surfaces, such as process catalyst deactivation. This

gasification reaction is thermodynamically favored at high temperatures (T >

680oC). The reaction, if carried out alone, requires high temperature (for fast

reaction) and high pressure (fir high reactant concentrations) for significant

conversion. However, this reaction in practical gasification applications is almost

never attempted as a solo chemical reaction, because of a variety of factors


32

incluidng low conversion, slow kinetic rate, low thermal efficiency, unimpressive

process economics, etc.

Hydrogasification

Direct addition of hydrogen to coal under high pressure forms methane.

This reaction is called hydrogasification and may be written as.

C(s )+2H 2 (g)↔CH 4(g)∆ H 298o =−74.8 kJ /mol (2.3)

This reaction is exothermic and is thermodynamically favored at low

temperatures (T < 670°C), unlike both steam and CO2 gasification reactions.

However, at low temperatures, the reaction rate is inevitably too slow. Therefore,

high temperature is always required for kinetic reasons, which in turn requires

high pressure of hydrogen, which is also preferred from equilibrium

considerations. This reaction can be catalyzed by K2CO3, nickel, iron chlorides,

iron sulfates, etc. However, use of catalyst in coal gasification suffers from

serious economic constraints because of the low raw material value, as well as

difficulty in recovering and reusing the catalyst. Therefore, catalytic coal

gasification has not been practiced much.

Partial Oxidation

Gasifying agent in partial oxidation is oxygen, which may be supplied as

pure oxygen or as air. Partial oxidation sometimes called as combustion. The

reactions for partial oxidation are shown as follows.

C(s )+12

O2 (g)

↔CO(g)∆ H 298o =−111.4 kJ /mol (2.4)

CO(g )+12

O2(g)

↔CO2(g)∆ H 298o =−393.5 kJ /mol (2.5)

Partial oxidation involves a complex reaction mechanism that depends on

how fast and efficiently combustion progresses. The reaction pathway is further

complicated because of the presence of both gas – phase homogenous reactions

and heterogeneous reaction between gaseous and solid reactans. Combustion or

oxidation of coal is much more complex in its nature than oxidation of carbon.

Coal is not a pure chemical species, rather it is a multifunctional,

multispecies, and heterogeneous macromolecule. It was found that in a region

where chemical reaction rate is controlling the overall rate. Typically in a low

temperature region where the kinetic rate is much slower than the diffusional rate


33

of reactant, the catalytic effect of mineral matter is a determining factor for coal

reactivity. It was also found that for high temperature regions where the external

mass transfer rate controls the overall rate, the reactivity of coal decreased with

increasing coal rank.

From the explanation, we can combine partial oxidation with steam

gasification. Partial oxidation can increase the reactivity of coal because our raw

material is low rank coal. Partial oxidation produce carbon monoxide and steam

gasification produce hydrogen. Besides that, both reactions can complete each

other in heat. Partial oxidation is exothermic process that would produce heat

from its reaction. Heat will be used for steam gasifying because its reaction is

endothermic and require heat.

b. Selection of Gasification Technology

Based on the reactor configuration. As weel as by the method of contacting

gaseous and solid streams, gasification process can also be categorized into the

following 3 types :

Fixed Bed Gasifier

It also called as fixed bed reactor because solids in the bed stay together

regardless of the movement of the hardware that supports of the bed. Coal is

supported by a grate and the gasifying media (steam, air, or oxygen) pass upward

through the supported bed, whereby the product gases exit from the top of the

reactor in the fixed bed reactor. Only noncaking coals can be used in the fixed bed

reactor. Coal and gaseous streams move counter currently. The temperature at the

bottom of rreactor is higher than that at the top. Because of the lower temperature

at the top for coal devolatiliation,relatively large amounts of liquid hydrocarbons

area also produced in this type of gasifier.


34

Figure 2.4. Fixed Bed Gasifier(Source : ETSAP, 2010)

The residence time of the coal is much longer than that in a suspension

reactor, thus providing ample contact time between reactants. Ash s removed from

the bottom of reactor as dry ash or slag. The example of this type of reactor are

Lurgi and Wellman – Galusha gasifiers.

Fluidized Bed Gasifier

It uses finely pulverized coal particles. The gas (or gasifying medium) flows

upward through the bed and fluidizes the coal particles. Owing to the ascent of

particles and fluidizing gas, larger coal surface area is made available, which

positively promotes the gas-solid chemical reaction, which in turn results in

enhancement in carbon conversion.


35

Figure 2.5. Fluidized Bed Gasifier(Source : ETSAP, 2010)

This type of reactor allows intimate contact between gas and solid coal

fines, at the same time providing relatively longer residence times than entrained

flow reactor. Dry ash is either removed continuously from the bed, or the gasifier

is operated at such a high temperature that it can be removed as agglomerates.

Such beds, however, have limited ability to handle caking coals, owing to

operational complications in fluidization characteristics. Winkler and Synthane

processes use this type of reactor.

Entrained Bed Gasifier

In this reactor, there is no bed of solids. This reactor system uses finely

pulverized coal particles blown into the gas stream before entry into the reactor,

with combustion and gasification occuring inside the coal particles suspended in

the gas phase. Because of the entrainment requirement, high space velocity of gas

stream and fine powdery coal particles are very essential to the operation of this

type of process. Because of the very short residence time (i.e., high space

velocity) in the reactor, a very high temperature is required to achieve good

conversion in such a short period of reaction time. This can also be assisted by

using excess oxygen.


36

Figure 2.6. Entrained Bed Gasifier(Source : ETSAP, 2010)

This bed configuration is typically capable of handling both caking and

noncaking coals without much operational difficulty. Examples of commercial

gasifiers that use this type of reactor include the Koppers-Totzek gasifier and

Texaco gasifier.

In the sub bab selecting gasification technology, we would select gasifier

that we used. In the previously paragraph, we have been already explained that

gasifier technology that we selected depends on various factor.

Conversion

Conversion always become one of the factor in selecting reactor, because it

clarify how many carbon in coal can reacts with gasification media (oxygen and

steam).

Total of gasification media

Gasification media that we used are steam and oxygen. Total of gasification

media that we used will be effected in utility cost.

Feedstock (coal rank)

Definition of feedstock (coal rank) is type of coal as input. Because in

Indonesia, there are so many varieties of coal.


37

Purity of Product

Product of the gasification process is synthetic gas. Purity of synthetic gas

means that pure from moisture.

Cost

Cost is investation cost from reactor unit that we used.

Priority grade at the table 2.1. are feedstock (coal rank), cost (OPEX), purity

of product (free from tars and oils), conversion, and total of gasification media.

Coal is used in the gasification will be variative appropriate condition in

Indonesia. Flexibility capability of reactor to process all variant of coal as feed

and stable product is important. So, feedstock become first factor to select this

grasifier. Second, cost is investation cost where it will be influence to economic

analysis. Third factor is purity of product. It is important because syngas will

become as raw material in synthesis methanol process. Higher purity of product

affected in easy treatment and more economic. The second last of the factor are

conversion and total of gasification media.

Table 2.1. Scoring the type of Gasifier

Gasifier

Type

Conversio

nFeedstock Cost

Total of Gasification

Media

Purity of

Product

Fixed Bed 5 1 3 5 1

Fluidized Bed 2 5 5 3 3

Entrained

Bed3 2 2 2 5

(Source : Author’s Personal Data)

In the table 2.1, we show that each gasifier has advantage and disadvantage,

such as fixed bed has advantage in high conversion and total of gasification media

is small, fluidized bed has advantage in feedstock (coal rank) because all variant

of coal capable become feed, and entrained bed has advantage in high purity of

product.

From scoring in the table 2.1., we choose gasifier that will be used is

Entrained Bed Reactor. This type of gasifier is choosen because high purity of

product.


38

2.3.2. Synthesis EthyleneFrom three methods option in section 2.1.3, we should to make scoring

process to select the best method to implement for the plant synthesis etyhelene

unit. In this scoring process, we use several parameters, such as how much

equipment needed, cost, raw material and the complexity the process. The

following table shows the score of available method. From this step, we will have

one method for produce ethylene. The process selection for produce ethylene will

explain in the following table:

Table 2.2. Scoring Synthesis Ethylene

Method

sEquipment Needed Raw Material

Cos

t

Complexit

yTotal

1 4 4 4 4 16

2 3 4 3 4 14

3 1 4 1 2 8


Raw Material

All of three methods have same raw material. From process before will

produce synthesis gas. So the score is same.

Equipment Needed

From all of three methods, methods one has the best score. Method one us

the simplest process. Method two is fermentation process and actually use batch

reactor. The process use many batch reactor, so the process can be called

contiunes (semi-batch).

Method three has a lot of equipment because the synthesis gas will be

convert to synthesis crude oil and then will reacted to produce ethylene.Thath

process need many equipment and unit, such as fischer-tropsch, treatment,

distillation section and ethylene production section.

Cost

Based on the amount of total equipment needed, more equipment means

more cost to spend to buy those equipment and also to maintain those equipment.

So, we can assume that more equipment nedded will cause higher cost for capital


39

cost and operational cost. Method three use fischer-tropsch which is expensive

process and use many catalyst for the process until produce the ethylene.

Complexity

Complexity is how easy the process to react or convert raw matrial/feed to

become the final product. From the equipment needed and process scheme, we

can see that method 1 is the highes complexity. Method 1 is the easiest one

because the process is relatively fast and not needed many process treatment.

Method two is almost same with method one, but to process the ethanol should be

procced by fermentation which is need long time enough. Method three has many

path of process because synthesis gas will be convert to synthesis crude oil which

has many treatment to forwarded to the next process. Method three also use

thermal cracking which has long and need a lot of equipment process.

Complexity means how easy the process to convert raw material to be

ethylene. It also related to the much amount of equipment involve on the process.

So, the best for this parameter is method one.

2.4. Process Description

2.4.1. Block Flow DiagramThe coal to ethylene process used for this project typically incorporates the

following process oerations, shown in Figure 2.7.


40

Figure 2.7. Block Flow Diagram for This Plant(Source : Author’s Personal Data)


41

2.4.2. Process Flow Diagram

Figure 2.8. Process Flow Diagram for Gasification Unit(Source : Author’s Personal Data)


42

Figure 2.9. Process Flow Diagram for Acid Gas Removal Unit(Source : Author’s Personal Data)


43

Figure 2.10. Process Flow Diagram for Water Gas Shift Unit(Source : Author’s Personal Data)


44

Figure 2.11. Process Flow Diagram for Synthesis Methanol Unit(Source : Author’s Personal Data)


45

Figure 2.12. Process Flow Diagram for Synthesis Ethylene Unit(Source : Author’s Personal Data)


46

2.4.3. Process Description

2.4.3.1. Gasification TechnologyCoal is pre-processed by methods such as crushing, sizing and drying to

prepare it for the coal gasification process. Coal is crushed into fine particles until

the diameter of coal is less than 100 mesh. Crushed coal feedstock combining

with water in a tank would make slurry. We have to control the water content of

coal less than 20% feed, so that the coal doesn’t absorb much water. In process,

we use pure water in making slurry to avoid impurities in water that can hamper

the process. The criteria of coal slurry in gasification process are homogeneous

coal size, low viscosity, and contain high coal carbon. The slurry is introduced

with oxygen through the feed-injector (burner) into the refractory-lined gasifier.

Then crushed coal is entering conversion process. During the coal

conversion process, the coal is decomposed in a high pressure and temperature

using steam and an oxygen supply, represented by the following equations.

In gasifying agent selection, steam and oxygen was chosen. Gasifying

agent reacts with coal in partial oxidation combine with steam gasification. First,

reaction will occur in exothermic. Then, the heat from that reaction would be

using in endothermic for the next reaction. In this process will produce raw

syngas. In gasifiers, carbon in coal slurry is converted to syngas, and the mineral

matters are transformed to ash/slag. The majority of the ash is melted and

deposited on the walls of the gasifier, forming a liquid slag, which flows out of

the bottom of the gasifier and finally solidifies in a water bath. However, a small

fraction of the ash is entrained as fly ash with the raw syngas out of the gasifier to

downstream processing.

Our plant use entrained flow reactor for gasification unit. Entrained bed

gasifier is horizontal gasifier operates at atmospheric pressure or a little higher

than atmospheric pressure. If this gasifier operates at high pressure, so the results

are a few tars and oil in gas product will be made. This gasifier can operates at

low pressure to maintain ash, so it will be dry solid. This gasifier can also operates

at temperatures above melting point of ash, so it will be liquid melted.

The shape of raw material in entrained bed gasifier must be small (< 0.1

mm) and homogen. The fuel will be entered into gasifier together with


47

gasification medium at high flowrate (oxygen and steam) in specific operation so

it can made particles. The fuel and gasification medium flow in the same direction

(co – current). The shape of material pull through approximately 80 % less than

200 mesh (44 µm). The residence time of this gasifier is about 1 – 10 seconds.

Because of the short residence time, entrained flow gasifiers must operate at high

temperatures to achieve high carbon conversion of about 99%. Since the gasifier

operates at a high temperature, the destruction of tar and oil yields a very pure

syngas. Operation temperature of this gasifier is high, more than 1000oC and

uniform in the entire side of gasifier in order that more producer gas is obtained

and producer gas is not bind tar and methane.

Ash is taken as slag because operation temperature more than melting

point of ash. Output temperature of product at this type of gasifier approximately

900 – 1400oC. Control at this gasifier are flowrate of the fuel, oxygen, and steam.

Efficiency of gasifier will be determined by operation temperature, particle size,

and injection rate of steam.

The advantage of this type gasifier are not too observe characteristics of

raw material, suitable for small raw material, producer gas involving a few of tar,

ash is taken as slag, product at high temperature need quenching to cleaning,

cooling can do by reuse so the result is more efficient.

The gasification process occurs generally around 1500 ºC, which is in the

slagging temperature range (Sheida et al, 2008). The pressure requirement is

between 30-70 bar. The produced syngas is cooled in a water quench because it is

suitable for providing conditions for CO shift reaction, required in the coal-to-

methanol process. In using the quench process for syngas cooling, the gaseous

effluent leaves through the bottom of the reactor with the liquid ash and enters the

cooling vessel.

Molten/sticky ash in the fly ash could cause fouling of the syngas cooler.

One method to reduce syngas cooler fouling is to maximize slag and minimize fly

ash. For ash/intermediate char-slag particles to be trapped in the slag layer and not

rebound (elastic reflection), particle surface stickiness and slag surface stickiness

are critical among many factors, such as particle velocity, surface tension,


48

temperature, size, and impact angle. Particles with certain carbon conversion at

temperature above the ash fusion temperature are sticky.

2.4.3.2. Acid Gas Removal

Raw Syngas from gasification unit contain impurities such as H2S.

Impurities limit applicability of syngas to chemical synthesis, so this plant need

acid gas removal unit.

Physical solvents tend to be favored over chemical solvents when the

concentration of acid gases or other impurities is very high. Unlike chemical

solvents, physical solvents are non-corrosive, requiring only carbon steel

construction.

At low partial pressures, physical solvents are impractical because the

compression of the gas for physical absorption is expensive. However, if the gas

is available at high pressure, physical solvents might be a better choice than

chemical solvents.

From the explanation above, physical solvent is the best choice for acid

gas removal method. Compare with another method, physical solvent have the

cheapest cost because it don’t need increase feed pressure or special material for

service. Physical solvent only need carbon steel construction.

A number of physical solvents are available for use in acid gas treating

processes. Four of the solvents are considered here: Dimethyl Ether of

Polyethylene Glycol (DEPG), Propylene Carbonate (PC), N-Methyl-2-

Pyrrolidone (NMP), and Methanol (MeOH).

DEPG is a mixture of dimethyl ethers of polyethylene glycol

(CH3O(C2H4O)nCH3 (n is between 2 and 9) used to physically absorb H2S, CO2,

and mercaptans from gas streams. DEPG can be used for selective H2S removal

which requires stripping, vacuum stripping, or a reboiler. The process can be

configured to yield both a rich H2S feed to the Claus unit as well as bulk CO2

removal. Selective H2S removal with deep CO2 removal usually requires a two-

stage process with two absorption and regeneration columns. H2S is selectively

removed in the first column by a lean solvent that has been thoroughly stripped

with steam, while CO2 is removed in the second absorber.


49

The second stage solvent can be regenerated with air or nitrogen for deep

CO2 removal, or using a series of flashes if bulk CO2 removal is required. DEPG

also dehydrates the gas and removes HCN. Compared to the other solvents, DEPG

has a higher viscosity which reduces mass transfer rates and tray efficiencies and

increases packing or tray requirements, especially at reduced temperatures. Since

it is sometimes necessary to reduce temperature to increase acid gas solubility and

reduce circulation rate, this could be a disadvantage. DEPG requires no water

wash to recover solvent due to very low vapor pressure. DEPG is suitable for

operation at temperatures up to 347°F (175°C). The minimum operating

temperature is usually 0°F (-18°C).

Selexol was the chosen solvent for acid gas removal. Among the other

solvents, DEPG give the minimum cost because DEPG don’t need water wash to

recovery or refrigerant for its process. Besides that, selexol can remove H2S, CO2,

mercaptans and HCN from syngas.

The inlet raw syngas is contacted counter current in packed bed with

Selexol solvent. The temperature of syngas gets reduced in absorber. The amount

of H2S removed from the gas stream is about 95% of the total H2S contained.

Column is maintained around 150°C and 6 MPa.

After the acid gas removal process, the acid gas produced is usually

emitted in the form of H2S which tends to be highly toxic. It is necessary to

convert this H2S gas into two alternative products of sulfur using additional

process.

2.4.3.3. Water Gas Shift Clean syngas that contain low impurities continue to water gas shift

process. The governing equation for water-gas shift reactor is as follows.

Water gas shift is to adjust CO/H2 ratio of gas become 1:2. Steam is

added to CO-H2 feed mixture before entering the water-gas shift reactor to convert

CO to CO2 and additional H2. The water-gas shift reaction is commonly run at low

temperature since thermodynamic equilibrium favors high conversion of CO and

steam to CO2 and H2 in the presence of catalysts that enhance the reaction rate.

Although the equilibrium favors formation of product at lower temperatures, the

reaction kinetics is faster at elevated temperatures.


50

Because of that, water gas shift has two stages. The first step involves a

high temperature step operating at 310-450°C, which reduces CO content to 2-

3%. The second stage uses a low temperature water gas shift catalyst operates at

210–240°C. Most of the CO shift reaction out in industry, are converting 90% CO

to H2 in the first HTS reactor, and the 90% of the remaining CO is converted in

LTS reactor. The exit gas of the HTS unit must be cooled. This is usually done by

quenching with water, thus providing additional steam to the process.

2.4.4.4. Synthesis Ethylenea. Overall Process Description

Syngas to Methanol

The overall process that occurs at methanol plant consists of four stages,

which is syngas compression, methanol reactor and distillation section. H2S free

syngas from gasification unit is compressed from 20 bar to 43 bar. After

compression, syngas is then going through gas/gas exchanger to raise the gas’

temperature before entering the reactor. Syngas is cross-exchanged with the hot

gas exiting the methanol reactor until it reaches the temperature of 235°C.

Methanol reactor is a shell and tube reactor, which the catalysts are packed

in tubes. The catalyst used in this reactor is copper-based S 3-86 catalyst. The

main reactions of methanol synthesis which are mentioned before take place in

this reactor. Besides producing methanol, this reactor is also producing several

byproducts, such as dimethyl ether, methyl formate, ethanol and isobutanol. The

production of byproduct in the reactor depends on CO2/CO ratio, feed’s purity and

catalyst’s age. This reactor is kept in isothermal state, because isothermal reactor

produces fewer byproducts, provides high heat reaction recovery and easier

temperature control. Generated heat from the isothermal reaction in methanol

synthesis is used to heat water in the boiler. The steam generated in the boiler is

used in steam separator.

Methanol exiting the reactor is now about 260°C in temperature and 4%

methanol by volume. The exiting stream is then cross-exchanged with the feed

stream. The product of methanol reactor is then transferred into vertical separator

and crude methanol tank, before entering the distillation section. Vertical

separator separates raw methanol and purge gas. Raw methanol will enter the next


51

process, while purge gas will be compressed before circulated again to methanol

reactor with fresh steam. After being separated in vertical separator, methanol is

stored temporarily in crude methanol tank. Methanol is also added by NaOH 1-

2% for neutralizing the acids produced by side reactions in methanol reactor.

Since methanol content in the output of methanol reactor is only about 4%

weight, methanol needs to be purified. Raw methanol is purified in distillation

section until it reaches AA grade. The steam needed for distillation process is

provided by boiler which is heated by the heat generated in methanol reactor. The

distillation section consists of two stages, topping section and refining section.

Topping section separates methanol and heavier alcohol with impurities such as

CO, CO2, H2, N2, CH4, aldehyde, ketone and dimethyl ether, while refining section

separates methanol and heavier alcohol. The product of this section is AA grade

methanol.

Methanol To Olefins

Methanol to olefin synthesis process is a commercially valuable process,

particularly because of the high-demand for ethylene. Today these compounds are

created mainly through the non-catalytic cracking of naphtha under steam. The

methanol to olefin (MTO) process, however, uses a molecular-sieve catalyst that

efficiently converts methanol into olefins (propylene and ethylene). The

molecular-sieve under consideration is silicon, aluminum, phosphate and oxygen

based, and is hence called “SAPO-34”.

The mechanisms MTO leading to the production of large organic

compounds inside the pores of the catalyst, the role of these organic compounds in

the formation of propylene and ethylene, the discrepancy between observed

product ratios of propylene and ethylene and corresponding thermodynamic

predictions, and finally, the reactivity of the products under MTO conditions.

Large organic compounds are typically N-methyl-benzenes formed after a kinetic

induction period, and only in the presence of impurities in the methanol feed. The

successive methlyation of N-methyl-benzene is followed by de-ethylation and de-

propylation to produce ethylene and propylene respectively, the less methyl

substituted the benzene is, the higher the selectivity for ethylene. Under typical

reaction temperatures (~650 K), thermodynamics predicts that the ratio of


52

propylene to ethylene is between 0.5-1. Feeding propylene back into the catalyst

lowers selectivity for propylene, indicating that it can act as both a product and a

reactant. The same holds for ethylene.

MTO process converts methanol to ethylene and propylene at above 80%

carbon selectivity in a fluid bed reactor with continuous regeneration. Because of

the quick deactivation of MTO catalyst, a kind of high efficiency fast fluidized

bed reactor is adopted, the activity of deactivated catalyst is recovered by burning

the coke in the regenerator. Coke accumulates on the catalyst must be removed to

maintained catakyst activity. The catalyst is maintained by continuous transfer of

coked catakyst from reactor to the regenerator where the coke is burned by air.

The carbon selectivity apporaches 90% if the heavier fraction such as butenes are

also accounted for as part of the product. From this process will produce co-

product include very small amounts of C1-C4 parrafins, hydrogen, CO and CO2, as

well as ppm levels of havier oxygenates.

Methanol will enter the MTO reactor to react for produce olefins, the

reaction is :

2CH3OH C2H4 + 2H2O

3CH3CH2OH C3H6 + 3H2O

MTO reaction is conversion of alcohol (methanol) into a light alkene

(ethylene and propylene). To increase the selectivity ethylene, the reaction can be

added by water.

The reaction will produce water, so the stream which is output from reactor

will be continue to water separation and then will compressed to light olefin

recovery. From this section the ethylene and propylene will be separated until the

grade that we want. The detail of process will be explained in the next section.

b. Process Description On The Equipment

Syngas to Methanol

1. Compressor

Compressor’s function is to compress gases exiting from the acid gas

removal unit from 5170 kPa into 27500 kPa. This compressor is driven by electric

motor.


53

2. Gas/gas Exchanger

Gas/gas exchanger is needed to raise syngas’ temperature before entering

methanol reactor. Syngas’ temperature needs to be adjusted to match with the

catalyst of methanol reactor’s activity.

3. Methanol Reactor

Synthesis of syngas into ethanol takes place in this reactor. Heated syngas

feed will enter the tube side of the reactor, in which the catalysts are packed. This

reactor operates at temperature of 324,7°C and pressure of 25000 kPa.

4. Vertical Separator

Vertical separator separates raw methanol from byproducts produced in

methanol synthesis reactor. This separator operates at 10000 kPa and 75°C.

5. Crude Methanol Tank

Crude methanol tank stores raw methanol temporarily before fed into the

distillation process. This tank operates at 0,5 bar and 70°C.

6. Distillation Column

There are two distillation columns in this unit, which are the topping column

and refining column. Topping column is used to remove byproducts from

methanol and heavier alcohol fraction and operates at 3,5 bar and 130°C.

Meanwhile, refining column is used to purify methanol so it reaches AA grade.

Refining column operates at 40 bar and 150°C.

Methanol To Olefins

1. Reactor

The spesification of reactor R-101 is

Function : Place where the methanol conversion to olefins is held

Type : Fluidized Bed Reactor with Regenerator

Operation : P = 15 to 45 psig and T = 650 to 1000 OF

Characteristic : Conversion 80%-90%


54

Figure 2.13. Methanol to Olefins Reactor(Source : A, Grefory, et.al. Hydrocarbon Engineering)

In this reactor, methanol will be held to dehydration reaction with shape

selective transformation to low molecular weight alkenes, such as ethylene,

propylene and havier alkenes fraction and also impurities. The impurities from

this reaction is H2O, H2, COx, light paraffins, and coke.

The reactor in MTO reactor use particular molecular sieve catalyst, that is

SAPO-34. Sapo-34 is a porous mineral, a crystalline silicon aluminum phosphate

molecular sieve (zeoliteis a crystalline aluminum silicate).It is like aluminum-

phosphate with some of the phosphorous (P+5) atoms substituted by silicon (S+4)

atoms. SAPO-34 is a caalyst for the methanol to olefins reaction.

The catalyst will give direction of reaction to produce ethylene and

propylene. The reaction is not only produce the ethylene and propylene, but also

produce the impurities and also the intermediet product like mechanisms below :

Figure 2.14. Reaction to Produce Ethylene and Propylene(Source : A, Grefory, et.al. Hydrocarbon Engineering)


55

The yield of olefins increases with the number of methyl groups on benzene.

Increasing space velocity of methanol, decreased ethylene selectivity. Addition of

water increased ethylene selectivity When benzenes approximately have 2

methyls, then transition occurs for making more propylene.

From experiment, SAPO-34 is more selective towards ethylene at elevated

temperatures. Both correlations increase ethylene content with increasing

temperature because the reaction are endothermic. SAPO-34 has a higher molar

ratio.

Figure 2.15. Diagram of Temperature versus Selectivity(Source : A, Grefory, et.al. Hydrocarbon Engineering)

Selectivity for ethylene decreases with additional ethylene in the feed.

Indicates that ethylene is being consumed to form other products. It is called

ethylene to propylene. Selectivity for propylene decreases with additional

propylene in the feed. Indicates that propylene is being consumed to form other

products. It is called propylene to ethylene.


56

Figure 2.16. Feed Effect to Selectivity(Source : A, Grefory, et.al. Hydrocarbon Engineering)

2. Separation Train

Distillationis a process which is generally used to separate a mixture of two

or more liquids based on their boiling points. However, what happens in a

distillation column is essentially a series of flashes, which are connected with

recycle loops. The liquid from each tray comes to equilibrium (ideally) with the

vapor, and the vapor rises up to the next tray and the liquid falls to the tray

beneath it. Each tray has a different temperature because a reboiler on the bottom

and a condenser at the top maintain a temperature gradient across the column.

When it reaches very low temperatures, the gas mixture becomes a liquid.

The different components making up the gas are then separated by carefully

boiling them off in a process known as distillation (or fractionation).

The cryogenically chilled stream is processed through a series of distillation

columns. Several columns are needed to seperates out th desired product. This

process section consists a demethanizer, ethane/ethylene and propane/propylene

splitters.

In this separation section, the dried gases are cooled to -120 °C/-184 °F and

pass through the demethanizer. The C1 or methane cut is further separated to

recover methane, and the bottom product is sent to the de-ethaniser. The de-

ethaniser top product is sent through an ethanelethylene splitter.. The bottoms of

the de-ethaniser are sent to the depropaniser for separation of the propylene. The

bottoms of the depropaniser are sent to the will recycled to the process.


57

3. De-Ethaniser

The de-ethanizer column functions as a fractionator to separate out as much

ethane as possible. The liquid feed is introduced on the top tray of the de-ethaniser

column, which is heated and operating at 25-30 bar/360-435 psi pressure. An

asymmetrical separation is effected so that a very low concentration of ethane

exits by way of the bottom product, and some heavier components remain in the

top product.

4. Ethylene Fractionator

The de-ethaniser top product is usually sent to the ethylene fractionator,

where ethylene is separated at very high purity by cryogenic distillation from the

ethane. The product from Ethylene fractinator are 99.95% ethylene, 0.01%

methane and 0.04% ethane. The operation condition of ethylene fractionator are at

pressure 15 atm amd temperature -37.5 OC.

Table 2.3. Condition Operation in Each Equipment (Methanol to Olefin)

No Equipment Function Type AmountCondition Operation

1 De-ethaniserTo seperate ethane

contain from hydrocarbon feed

Sieve tray Plate Tower

1P = 30 atm

T= -36 - 135oC

2Ethylene

fractinator

To seperate ethylene contain

from hydrocarbon feed

Sieve tray Plate Tower

1P= 15 atm

T=-37.5 – (-16) oC



58

CHAPTER IIIMASS & ENERGY BALANCE

From the process in chapter II, we can calculate mass and energy balance of

the plant. Mass balance will be evaluated each equipment. From this chapter, we

will know how much coals needed to produce ethylene with our capacity, 400.000

tons per year. Then, we will know the energy needed to generate the process from

start until ethylene produced. In this chapter, mass and energy balance will be

calculated each unit and then each equipment.

3.1. Mass Balance for Equipment

3.1.1. Gasification Unit

Figure 3.1. The Hysys Simulation for Gasification Unit(Source : Author’s Personal Data)

Mixer (M – 100)

Table 3.1. show about the mass balance for the mixer M – 100 which contains

of C, H2, O2, S, N2 (stream 2) and H2O (stream 4) as the composition inlet and

outlet. In the table 3.1, we can see mass balance in mixer.

Table 3.1. Mass Balance on Mixer

Component

s

Input (ton/d) Output (ton/d)

2 4 5

C 3030.66. 0.00 3030.66

CO 0.00 0.00 0.00

CO2 0.00 0.00 0.00


59

Table 3.1. Mass Balance on Mixer (Cont’d)

Component

s


2 4 5

H2O 588.03 2063.85 2651.89

H2 212.59 0.00 212.59

O2 1031.32 0.00 1031.32

H2S 0.00 0.00 0.00

S 18.99 0.00 18.99

N2 23.97 0.00 23.97

TOTAL 4905.60 2063.85 6969.45


Slurry Pump (JC - 101)

Table 3.2. show about the mass balance for the slurry pump JC – 101 which

contains of C, H2, O2, S, N2, and H2O as the composition inlet and outlet. In the

table 3.2, we can see mass balance in slurry pump.

Table 3.2. Mass Balance on Slurry Pump

Component

s


5 6

C 3030.66 3030.66

CO 0.00 0.00

CO2 0.00 0.00

H2O 2629.82 2629.82

H2 3.23 3.23

O2 33.19 33.19

H2S 0.00 0.00

S 18.99 18.99

N2 0.84 0.84

TOTAL 5679.85 5679.85


Gasifier (R – 101)


60

The type of gasifier that we use in this plant is entrained flow gasifier. It

accepts almost any type of coal, including caking coal, without any major

operational restrictions. It has the highest operating temperature (around 1400–

1500°C) of all the conventional gasifiers. The process has a track record of over

50 years of safe operation. The overall thermal efficiency of the gasifier is 85 to

90%. The time on stream (TOS) or availability is better than 95%.

Then, table 3.3. is about the result of mass balance from gasifier which stream

8 (oxygen), stream 3 (water, hydrogen, oxygen, nitrogen), and stream 6 (carbon,

water, hydrogen, oxygen, sulphur, nitrogen) as the input and stream 16 and 10

(carbon monoxide, carbon dioxide, water, hydrogen, hydrogen sulphide, nitrogen)

as the ouput

Table 3.3. Mass Balance on Gasifier

ComponentsInput (ton/d) Output (ton/d)

8 3 6 16 10

C 0.00 0.00 3030.66 0.00 0.00

CO 0.00 0.00 0.00 0.00 5395.97

CO2 0.00 0.00 0.00 0.00 2622.91

H2O 0.00 22.07 2629.82 0.00 361.21

H2 0.00 212.56 3.23 0.00 467.74

O2 1926.49 1030.99 33.19 0.00 0.00

H2S 0.00 0.00 0.00 0.00 20.18

S 0.00 0.00 18.99 0.00 0.00

N2 0.00 23.96 0.84 0.00 23.97

TOTAL 1926.49 1289.60 5679.85 0.00 8892.01


Cooler Raw Syngas (E – 101)

Cooler raw syngas is an cooler which has function to cooling raw syngas after

gasification process and before enter to cyclone. Table 3.4. below is the table

of mass balance between stream 11(the result of gasification) and stream 11 as

the component whic willl be entered to the cyclon.

Table 3.4. Mass Balance on Cooler Raw Syngas


61


10 11

C 0.00 0.00

CO 5395.97 5395.97

CO2 2622.91 2622.91

H2O 361.21 361.21

H2 467.74 467.74

O2 0.00 0.00

H2S 20.18 20.18

S 0.00 0.00

N2 23.97 23.97

TOTAL 8892.01 8892.01


Cyclone Separator (FG 101)

Cyclone separator has function to separate raw syngas from slag/ash. Table 3.5.

shows about mass balance stream 11 as input and stream 12 and 15 as the

ouput.

Table 3.5. Mass Balance on Cyclone Separator


11 15 12

C 0.00 0.00 0.00

CO 5395.97 1.75 5395.96

CO2 2626.91 1.30

2625.6

1

H2O 361.21 339.86 21.35

H2 467.74 0.13 467.74

O2 0.00 0.00 0.00

H2S 20.18 3.32 20.15

S 0.00 0.00 0.00

N2 23.97 0.04 23.97

TOTAL 8896.01 341.22

8554.7

9


62


3.1.2. Acid Gas Removal Unit

Figure 3.2. shows about the hysis simulation for acid gas removal unit. Acid gas

removal is a step after gasification. The purpose of the process is to remove

undesired component from syngas, such as H2S or CO2

Figure 3.2. The Hysys Simulation for Acid Gas Removal Unit(Source : Author’s Personal Data)

Raw Syngas Compressor (JC – 201)

Table 3.6. show about the mass balance for the compressor which contains of

stream 12 as the composition inlet and stream 13 as composition outlet. In the

table 3.6, we can see mass balance in raw syngas compressor.

Table 3.6. Mass Balance on Raw Syngas Compressor

ComponentInput (ton/d) Output (ton/d)

12 13

CO 5370.60 5370.60

CO2 2654.79 2654.79

H2O 20.98 20.98

H2 464.15 464.15

H2S 19.82 19.82

N2 24.46 24.46DEPG 0.00 0.00

TOTAL 8554.79 8554.79(Source : Author’s Personal Data)


63

Raw Syngas Cooler (E – 201)

Table 3.7. below is the table of mass balance between stream 13 (input) and

stream 14 (output).

Table 3.7. Mass Balance on Raw Syngas Cooler


13 14

CO 5370.60 5370.60

CO2 2654.79 2654.79

H2O 20.98 20.98

H2 464.15 464.15

H2S 19.82 19.82

N2 24.46 24.46DEPG 0.00 0.00


Absorber Column (T – 201)

Absorber column is used to absorb raw syngas become clean syngas using

DEPG solvent. The type of absorption in this this process is amine absorption.

Then, table 3.8. is about the result of mass balance from absorber column

which stream 14 and 31 as the input and stream 15 and 32 as the ouput.

Table 3.8. Mass Balance on Absorber Column


14 31 32 15

CO 5370.60 0.00 5345.82 22.10CO2 2654.79 0.00 1901.84 751.62H2O 20.98 6008.55 0.92 6028.57H2 464.15 0.00 463.68 0.24

H2S 19.82 0.36 0.15 20.05N2 24.46 0.00 24.35 0.09

Table 3.8. Mass Balance on Absorber Column (Cont’d)


64


14 31 32 15

DEPG 0.00 50822.85 0.00 50822.85

TOTAL 8554.79 56831.75 7736.76 57645.53(Source : Author’s Personal Data)

Rich DEPG Pump (J – 201)

Rich DEPG Pump is used to flowing the DEPG solvent by increasing the

pressure of the liquid. In this case, the DEPG solvent is pumped to absorber

column to absorb undesired component. Table 3.9 shows about the mass

balance of the Rich DEPG Pump.

Table 3.9. Mass Balance on Rich DEPG Pump


15 16

CO 22.10 22.10CO2 751.62 751.62H2O 6028.57 6028.57H2 0.24 0.24

H2S 20.05 20.05N2 0.09 0.09

DEPG 50822.85 50822.85


Stripper Column (T – 202)

Stripper column is used to regenerate raw syngas. Then, table 3.10. is about the

result of mass balance from stripper column which stream 16 as the input and

stream 17 and 23 as the ouput

Table 3.10. Mass Balance on Stripper Column


65

Componen

t


16 17 23

CO 22.10 22.10 0.00CO2 751.62 751.62 0.00H2O 6028.57 21.68 6006.91H2 0.24 0.24 0.00

H2S 20.05 19.69 0.36N2 0.09 0.09 0.00

DEPG 50822.85 1.10 50821.75TOTAL 57645.53 816.48 56829.02


Rich DEPG Cooler (E – 203)

Rich DEPG Cooler is used to cooling the DEPG solvent. Table 3.11. shows

about the mass balance of the Rich DEPG Cooler which is stream 27 as the

input and stream 28 as the output.

Table 3.11. Mass Balance on RICH DEPG Cooler

Componen

t


27 28

CO 0.00 0.00CO2 0.00 0.00H2O 6006.91 6006.91H2 0.00 0.00

H2S 0.36 0.36N2 0.00 0.00

DEPG 50821.75 50821.75TOTAL 56829.02 56829.02


Rich DEPG Pump (J – 204)

Rich DEPG Pump is used to flowing the DEPG solvent by increasing the

pressure of the liquid. In this case, the DEPG solvent is pumped to absorber

column to absorb undesired component. Table 3.12 shows about the mass

balance of the Rich DEPG Pump.


66

Table 3.12. Mass Balance on Rich DEPG Pump

Componen

t


26 27

CO 0.00 0.00CO2 0.00 0.00H2O 6006.91 6006.91H2 0.00 0.00

H2S 0.36 0.36N2 0.00 0.00

DEPG 50821.75 50821.75TOTAL 56829.02 56829.02


3.1.3. Water Gas Shift UnitFigure 3.3. shows about the hysis simulation for water gas shift reaction unit.

Water gas shift reaction is a final step in gasification. The purpose of the process

is to increase ratio CO and H2.

Figure 3.3. The Hysys Simulation for Water Gas Shift Unit(Source : Author’s Personal Data)

Expander (K – 100)


67

Table 3.13. show about the mass balance for the expander which contains of

stream 21 as the composition inlet and stream 22 as the composition outlet. In

the table 3.13, we can see mass balance in slurry pump.

Table 3.13. Mass Balance on Expander


21 22

C 0 0

CO 5910.55 5910.55

CO2 2101.60 2101.60

H2O 92.59 92.59

H2 512.54 512.54

O2 0.00 0.00

C1 0.00 0.00

C2 0.00 0.00

C3 0.00 0.00

H2S 60.96 60.96

S 0.00 0.00

N2 27.35 27.35

C2H5OH 0.00 0.00

CH3OH 0.00 0.00

TOTAL 8553.60 8553.60


Reboiler (E – 106)

Table 3.14. show about the mass balance for the reboiler which contains of

stream 22 as the composition input and stream 29 as composition output

Table 3.14. Mass Balance on Reboiler

ComponentInput Output

22 29

C 0 0

CO 5910.55 5910.55

CO2 2101.60 2101.60

Table 3.14. Mass Balance on Reboiler (Cont’d)

Component Input Output


68

22 29

H2O 92.59 92.59

H2 512.54 512.54

O2 0.00 0.00

C1 0.00 0.00

C2 0.00 0.00

C3 0.00 0.00

H2S 60.96 60.96

S 0.00 0.00

N2 27.35 27.35

C2H5OH 0.00 0.00

CH3OH 0.00 0.00

TOTAL 8553.60 8553.60


High Temperature Shift (GBR – 102)

High temperature shift operating at 310 – 450°C, which reduces CO content to

2 – 3%. Table 3.15. show about the mass balance for the high temperature shift

which contains of stream 29 and stream 30 as the composition input and stream

31 and dummy 1 as composition output.

Table 3.15. Mass Balance on High Temperature Shift


29 30 31 DUMMY 1

C 0 0.00 0.00 0.00

CO 5910.55 0.00 5985.54 0.00

CO2 2101.60 0.00 1983.78 0.00

H2O 92.59 162.49 50.77 0.00

H2 512.54 0.00 507.14 0.00

O2 0.00 0.00 0.00 0.00

C1 0.00 0.00 0.00 0.00

Table 3.15. Mass Balance on High Temperature Shift (Cont’d)



69

29 30 31 DUMMY 1

C2 0.00 0.00 0.00 0.00

C3 0.00 0.00 0.00 0.00

H2S 60.96 0.00 0.60 0.00

S 0.00 0.00 0.00 0.00

N2 27.35 0.00 27.35 0.00

C2H5OH 0.00 0.00 0.00 0.00

CH3OH 0.00 0.00 0.00 0.00

TOTAL 8553.60 162.49 8555.22 0.00


Reboiler (E – 107)

Table 3.16. show about the mass balance for the reboiler which contains of

stream 31 as the composition input and stream 32 as composition output.

Table 3.16. Mass Balance on Reboiler


31 32

C 0.00 0.00

CO 5985.54 5985.54

CO2 1983.78 1983.78

H2O 50.77 50.77

H2 507.14 507.14

O2 0.00 0.00

C1 0.00 0.00

C2 0.00 0.00

C3 0.00 0.00

H2S 0.60 0.60

S 0.00 0.00

N2 27.35 27.35

C2H5OH 0.00 0.00

CH3OH 0.00 0.00

TOTAL 8555.22 8555.22


Low Temperature Shift (GBR – 103)


70

Low temperature water gas shift catalyst operates at 210 – 240°C. The 90% of

the remaining CO is converted in LTS reactor. Table 3.17. show about the

mass balance for the low temperature shift.

Table 3.17. Mass Balance on Low Temperature Shift

Componen

t

Input Output

32 33 34 DUMMY

C 0.00 0.00 0.00 0.00

CO 5985.54 0.00 5911.83 0.00

CO2 1983.78 0.00 2099.59 0.00

H2O 50.77 162.49 4.99 0.00

H2 507.14 0.00 512.45 0.00

O2 0.00 0.00 0.00 0.00

C1 0.00 0.00 0.00 0.00

C2 0.00 0.00 0.00 0.00

C3 0.00 0.00 0.00 0.00

H2S 0.60 0.00 0.60 0.00

S 0.00 0.00 0.00 0.00

N2 27.35 0.00 27.35 0.00

C2H5OH 0.00 0.00 0.00 0.00

CH3OH 0.00 0.00 0.00 0.00

TOTAL 8555.22 162.49 8556.84 0.00


Cooler (E – 108)

Before syngas enter to the next step, syngas is cooled first in cooler. Table

3.18. below is the table of mass balance in cooler

Table 3.18. Mass Balance on Cooler


34 35

C 0.00 0.00

CO 5911.83 5911.83

Table 3.18. Mass Balance on Cooler (Cont’d)



71

34 35

CO2 2099.59 2099.59

H2O 4.99 4.99

H2 512.45 512.45

O2 0.00 0.00

C1 0.00 0.00

C2 0.00 0.00

C3 0.00 0.00

H2S 0.60 0.60

S 0.00 0.00

N2 27.35 27.35

C2H5OH 0.00 0.00

CH3OH 0.00 0.00

TOTAL 8556.84 8556.84


3.1.4. Synthesis Methanol Unit

Figure 3.4. The Hysys Simulation for Water Gas Shift Unit(Source : Author’s Personal Data)

Methanol synthesis unit is the unit that converts syngas into methanol. To

fulfill the production capacity, this unit has to produce 4577.65 tonnes of


72

methanol per day. The mass balance of methanol synthesis unit can be

summarized in the table below:

Compressor (JC – 401)

Table 3.19. Mass Balance on Compressor


41 (tonne/day) 42 (tonne/day)

CH3OCH3 0 0

CO 3864.31 3864.31

Table 3.19. Mass Balance on Compressor (Cont’d)

Componen

t

Input Output

41 (tonne/day) 42 (tonne/day)

CO2 4056.62 4056.62

H2O 54.18 54.18

H2 556.03 556.03

CH3OH 0 0

N2 24.27 24.27

TOTAL 8555.441 8555.441


From the table above, we can see that the input and output mass flow of the

compressor is balanced. There is no change of component composition in this

equipment because compressor only raise the stream’s pressure but not changing

its composition

Heat Exchanger (E – 401)

This heat exchanger is needed to exchange heat from methanol reactor’s feed

(stream 42) with the output of methanol reactor (stream 44). The heated stream

42 is then named stream 43, while the cooled stream 44 is named stream 45.

Table 3.20. Mass Balance on Heat Exchanger



73

42 (tonne/day) 45

(tonne/day)

44

(tonne/day)

43

(tonne/day)

CH3OCH3 0 30651.21 28525.88 0

CO 3864.31 67357.35 70599.30 3864.31

CO2 4056.62 312262.93 312763.57 4056.62

H2O 54.18 1441.67 405.64 54.18

H2 556.03 1697.78 2233.24 556.03

CH3OH 0 2681.37 1564.80 0

N2 24.27 25732.42 25732.42 24.27

TOTAL 8555.441 441824.74 441824.85 8555.44


From the table above, we can see that the input and output stream’s mass flow is

balanced. This balance can be seen as stream 42 has the same mass flow as stream

43, while stream 44 has the same mass flow as stream 45. There is no change of

composition in this equipment because there is only heat exchange in this equipment

without reaction.

Reactor (R – 401)

Table 3.21. Mass Balance on Reactor

Component

Input Output

44

(tonne/day)

45 (tonne/day)

CH3OCH3 28525.88 28525.88

CO 70599.30 70599.30

CO2 312763.57 312763.57

H2O 405.64 405.64

H2 2233.24 2233.24

CH3OH 1564.80 1564.80

N2 25732.42 25732.42

TOTAL 441824.85 441824.85


From the table above, it can be seen that the input and output of this reactor is

balanced in the term of mass. The input and output stream has the same mass flow


74

although they had different composition. The different composition is caused by

the reaction occurs in this reactor.

Separator (FG – 401)

Table 3.22. Mass Balance on Separator

Componen

t

Input Output

43 (tonne/day) 62 (tonne/day) 46 (tonne/day)

CH3OCH3 28525.88 28617.96 8637.559

CO 70599.30 67301.00 239.3709

CO2 312763.57 311640.54 2643.997

H2O 405.64 358.19 4602.792

H2 2233.24 1697.39 1.648575

CH3OH 1564.80 1582.76 4667.079

N2 25732.42 25729.45 12.60311

TOTAL 441824.85 436927.31 20805.05


From the table above, we can see that the input and the output of this equipment is

already balanced in term of mass. The two output stream’s mass flow if added is

equal to the input’s mass flow. Composition changes can also seen in the table,

because there are component separated in this equipment

Distillation Column (T – 401)

Table 3.24. Mass Balance on Distillation Column

Componen

t

Input Output


CH3OCH3 8637.559 0 8637.558

CO 239.3709 0 239.3709

CO2 2643.997 0 2643.997

H2O 4602.792 4602.792 0

H2 1.648575 0 1.648575

CH3OH 4667.079 4666.816 0.264219

N2 12.60311 0 12.60311

TOTAL 20805.05 9269.607 11535.44


From the table above, we can conclude that the input and output’s mass flow is

already balanced. If we add the two output’s mass flow, we will obtain the same


75

amount as the input’s mass flow. The table also show the amount of impurities

such as dimethyl ether, carbon dioxide, carbon monoxide, hydrogen, and nitrogen

removed from the raw methanol.

Distillation Column (T – 402)

Table 3.25. Mass Balance on Distillation Column

Componen

t

Input Output


CH3OCH3 0 0 0

CO 0 0 0

CO2 0 0 0

H2O 4602.792 25.60812 4577.184

H2 0 0 0

CH3OH 4666.816 4666.349 0.466414

N2 0 0 0

TOTAL 9269.607 4691.957 4577.65


From the table above, we can see that the input’s mass flow is already balanced

with the output’s mass flow. The addition of the two output’s mass flow is the

same with the input’s mass flow. From the table, we also can see how much water

that is removed to produce methanol grade AA

3.1.5. Synthesis Ethylene Unit

Figure 3.5. The Hysys Simulation for Synthesis Ethylene Unit(Source : Author’s Personal Data)


76

Synthesis ethylene unit is a net of equipments to convert methanol to

ethylene. From the plants. we hope that 1095.89 tons per day. Mass and energy

balance will be calculated by hysys simulation. Hysys simulation will give the

information of mass flow and energy flow and then we can evaluate about mass

and energy balance. The mass and energy balance for synthesis ethylene unit will

be explain bellow.

The mass balance for this unit is :

MTO Reactor (R-501)

Table 3.26. Mass Balance On MTO Reactor

REAKTOR R-501

Component Feed (Ton/d) Product(Ton/d)

Methanol 4583.59 769.78

Ethylene 0.00 1401.16

Propylene 0.00 258.84

Buthene 0.00 1.37

Penthene 0.00 1.36

Ethane 0.00 3.66

Propane 0.00 0.00

Methane 0.00 1.93

Water 0.00 2139.89

Table 3.26. Mass Balance On MTO Reactor (Cont’d)

REAKTOR R-501

Component Feed (Ton/d) Product(Ton/d)

Formic Acid - 5.58

TOTAL 4583.59 4583.60


From the table above. we can see that the mass flow inlet to reactor is equal

wih outlet from reactor.

Water Separation (V-501)

Table 3.27. Mass Balance On Water Separator

Water Separator

Componet Inlet(Ton/d) Top(Ton/d) Bottom(Ton/d)


77

Methanol 769.78 300.45 469.32

Ethylene 1401.16 1401.15 0.0051

Propylene 258.84 258.84 0.0002

Buthene 1.37 1.37 0.00

Penthene 1.36 1.36 0.00

Ethane 3.66 3.66 0.00

Propane 0.00 0.00 0.00

Methane 1.93 1.93 0.00

Water 2139.83 463.00 1676.88

Formic Acid 5.58 6.00 0.14

TOTAL 4583.60 2437.38 2146.36


From the table above. the inlet stream to water separation wil be separated

according to the operation condition. The mass flow inlet is equal with sum of

mass flows from outlet ( Top and Bottom).

Demethanizer (T-501)

Table 3.28. Mass Balance On Demethanizer

Demethanizer

Component Inlet(Ton/d) Top(Ton/d) Bottom(Ton/d)

Methane 1.93 1.93 0.0003

Table 3.28. Mass Balance On Demethanizer (Cont’d)

Demethanizer


Ethane 3.66 3.66 0.0002

Propane 0 0 0

Ethylene 1401.15 1401.00 0.15

Propylene 258.84 258.83 0.0098

Butene 1.37 1.37 0.0000

Pentene 1.36 1.36 0.0000

Water 463.00 19.57 443.42

Hydrogen 0 0 0

Methanol 300 133.68 166.77


78

Formic Acid 5.43 5.04 0.39

TOTAL 2437.22 1826.47 610.46


From the table above. we can see that the inlet will be distillated by

demethanizer to seperate stream inlet to result rich methane stream. The mass

flow of inlet is equal as with sum of outlet mass flow (top and bottom).

Deethanizer (T-502)

Table 3.29. Mass Balance On Deethanizer

Deethanizer


Methane 1.93 1.93 0

Ethane 3.66 3.66 0

Propane 0 0 0

Ethylene 1401.00 1400.99 0.00

Propylene 258.83 258.63 0.20

Butene 1.37 0.72 0.64

Pentene 1.36 0.04 1.32

Water 19.57 0.12 19.45

Hydrogen 0 0 0

Table 3.29. Mass Balance On Deethanizer (Cont’d)

Deethanizer


Methanol 134 0.0004 133.68

Formic Acid 5.04 0.0000 5.04

TOTAL 1826.48 1666.12 160.36


From the table above. we can see that the inlet will be distillated by

deethanizer to seperate stream inlet to result rich ethane and ethylene stream. The

mass flow of inlet is equal as with sum of outlet mass flow (top and bottom).

Ethylene Tower (T-503)

Table 3.30. Mass Balance On Ethylne Tower


79

Ethylene Tower


Methane 1.93 1.79 0.14

Ethane 3.66 2.53 1.19

Propane 0 0.0000 0.0000

Ethylene 1400.99 1364.65 36.33

Propylene 258.63 9.27 249.35

Butene 0.72 0.0000 0.72

Pentene 0.047 0.0000 0.047

Water 0.12 0.10 0.019

Hydrogen 0 0.0000 0.0000

Methanol 0.0004 0.0000 0.0004

Formic Acid 0.0000 0.0000 0.0000

TOTAL 1666.12 1378.07 288.05


From the table above. we can see that the inlet will be distillated by ethylene

tower to seperate stream inlet to result rich ethylene stream. The mass flow of

inlet is equal as with sum of outlet mass flow (top and bottom)

So from the last unit. our plant can produce the final product with ethylene purity

is 99% and side product that propylene with 80%.

3.2. Energy Balance for Equipment

3.2.1. Gasification UnitTable 3.31. Gasification Unit Energy Balance

Equipmen

tIn (x1014 kJ/h) Out (x1014 kJ/h) Require (x1014kJ/h)

M-101 2.000E+00 -3.6375.000E+0

0-1.708

4.000E+00 -1.349

-4.986 -1.708 -3.279

J-101 5.000E+00 -1.7086.000E+0

0-1.708


80

Q7 2.458

7.497 -1.708 2.457

R-101 3.000E+00 -5.0751.000E+0

1-1.713

6.000E+00 -1.7081.600E+0

10.00

8.000E+00 -7.764

-1.455E+15 -1.713 -1.283E+15

E-101 1.000E+01 -1.7131.100E+0

1-2.098

Q8 3.848

-1.713 1.750 -3.463

FG-101 1.100E+01 -2.0981.200E+0

1-1.874

1.500E+0

1-2.245

-2.098 -4.119 2.021

Overall Duty (kJ/h) -15.1


3.2.2. Acid Gas Removal UnitTable 3.32. Acid Gas Removal Unit Energy Balance

Equipment In (x1014 kJ/h) Out (x1014 kJ/h) Require (x1014 kJ/h)

JC-201 1.200 -6.328 1.300 -6.199

Q8 1.296

-5.033 -6.199 1.166

E-201 1.300E+01 -6.199 1.400E+01 -6.373

Table 3.32. Acid Gas Removal Unit Energy Balance (Cont’d)

Equipmen

tIn (x1014 kJ/h) Out (x1014kJ/h)

Require (x1014

kJ/h)

Q9 1.744

-6.199 -4.629 -1.570


81

T-201 3.100E+01 -1.502 1.500E+01 -1.601

1.400E+01 -6.373 3.200E+01 -5.390

-7.875 -6.991 -8.848

J-201 1.500E+01 -1.601 1.600E+01 -1.600

Q10 4.823

3.223 -1.600 4.823

T-202 1.600E+01 -1.600 1.900E+01 -9.889

Q10 4.823 2.600E+01 1.380

Q11 -5.681

3.223 -1.419E+15 1.741E+15

J-204 2.600E+01 1.380 2.700E+01 -1.373

Q12 7.537

8.917 -1.373 1.029E+15

E-203 2.700E+01 -1.373 Q13 1.290

2.800E+01 -1.502

-1.373 -2.121 -1.161

Overall Duty (kJ/h) 30.1


3.2.3. Water Gas Shift UnitTable 3.33. Water Gas Shift Unit Energy Balance

Equipmen

t In (x1014kJ/h) Out (x1014kJ/h) Require

JC301 32 -1.769 33 -1.785

Q 1.561

-1.769 -2.240 -1.545

E301 33 -1.785 34 -1.530

Q14 2.549

7.641 -1.530 2.294

Table 3.33. Water Gas Shift Unit Energy Balance (Cont’d)

Equipmen

t In (x1014 kJ/h) Out (x1014kJ/h) Require

R301 34 -1.530 38 -1.531


82

36 -8.866DUMMY

10.00

-1.040E+15 -1.531 -8.865

E302 38 -1.531 39 -1.641

Q15 1.099

-1.531 -5.421 -9.890

R302 37 -8.867 40 -1.642

39 -1.641 DUMMY 0.00

-1.051E+15 -1.642 -8.866

E303 40 -1.642 41 -1.680

Q16 3.782

-1.642 2.103 -3.744

Overall Duty (kJ/h) -21.72


3.2.4. Synthesis Methanol UnitEnergy balance of Methanol Synthesis Unit can be seen on the table below :

Table 3.34. Methanol Synthesis Energy Balance

Equipmen

tInput (kJ/h) Output (kJ/h) Duty (kJ/h)

JC-40141 42

-4.54E+10 -4.54E+10 0.00E+00

E-401

42 44

-4.54E+10 -4.38E+10 -1.57E+09

45 43

-4.38E+10 -4.04E+10 -3.43E+09

R-40144 45

-4.38E+10 -4.38E+10 0.00E+00

FG-401 43 46

-4.04E+10 -5.64E+10 1.60E+10

T-401 47 60 49

Table 3.34. Methanol Synthesis Energy Balance (Cont’d)

Equipmen Input (kJ/h) Output (kJ/h) Duty (kJ/h)


83

t

T-401 -5.64E+10 -1.84E+09 -1.73E+09 -5.28E+10

E-40650 48

-1.73E+09 -1.20E+09 -5.37E+08

E-40360 59

-1.84E+09 -2.09E+09 2.56E+08

T-40251 53 57

-3.47E+08 -4.54E+08 -7.62E+08 8.69E+08

E-40454 52

-4.54E+08 -4.12E+08 -4.19E+07

E-40757 58

-7.62E+08 -4.57E+08 -3.05E+08

Overall Duty (kJ/h) -4.16E+10


3.2.5. Synthesis Ethylene UnitTable 3.35. Energy Balance on Synthesis Ethylene

No Equipment

Inlet (x107

kJ/h)

Outlet (x107

kJ/h)

Duty (x107

kJ/h) State

1 R-501 -2.55 -2.55 -0.24 Balance

2 E-501 -2.55 -2.66 0.11 Balance

3 V-501 -2.66 -3.41 7.52 Balance

4 K-501 -5.04 -4.77 -2.69 Balance

5 E-502 -4.77 -5.48 0.71 Balance

6 T-501 -5.48 -6.16 0.67 Balance

7 K-502 1.69 1.67 0.24 Balance

8 T-502 1.67 1.25 0.43 Balance

9 K-503 2.48 2.44 4.30 Balance

10 E-503 2.44 2.58 0.18 Balance

11 T-503 2.58 2.30 0.279 Balance



84

From table above. if we sum the inlet column and sum outlet coulmn which

added with duty column. The inlet column will has same energy with outlet and

duty column.

3.3. Overall Mass BalanceFrom process flow diagram in previous chapter, we can calculate mass and

energy balance for this plant using Aspen HYSYS V7.3.

Table 3.36. Mass Balance Overall

Component

InletMass Flow ( Ton/d) Component outlet Mass Flow (ton/d)

C 3030.66 C 222.31

CO 0 Methane 79.2271

CO2 0 Ethane 149.991

H2O 588.03 Propane 0

H2 212.59 Ethylene 1451

O2 1031.32 Propylene 105.88

S 18.99 Butene 5.61

N2 23.97 Pentene 5.59

Hydrogen 0

Methanol 3.14

Formic Acid 2.28

S 18.99

N2 23.9739

Total 4905.60 Total 4905.60


3.4. Overall Energy BalanceTable 3.37. Energy Balance Overall

Component

InletEnergy Flow ( kJ/h)

Component

outletEnergy Flow (kJ/h)

C -87068 Methane -4681250

CO -4259301 Ethane -2824600

CO2 -9173407 Propane -2361136

Table 3.37. Energy Balance Overall (Cont’d)


85

Component

InletEnergy Flow ( kJ/h)

Component

outletEnergy Flow (kJ/h)

H2O -13984027 Ethylene 1868921

H2 -4143973 Propylene 486405

O2 -270609 Butene -124818

H2S -886442 Pentene -401271

S 8503897 Water -13434111

N2 -304632 Methanol -6290313

Formic Acid -8236739

H2S -886442

S 8503897

N2 -304632

Total -24605564 Total -24605564

(Source : Author Personal Data)

From the table above, we can see that the total of inlet mass and inlet energy

is will be equal with mass outlet and energy outlet.

3.5. Mass EfficiencyFrom mass balance calculation above, we can calculate mass efficiency for

main product produced in this plant. It can be done by dividing amount main

product produced with raw material used.

mass efficiency ethylene= et h ylene producedFeed Coal

= 14513030.66

x 100 %=45.47 %

mass efficiency Total= et h ylene producedFeed Coal

= 21253030.66

x100 %=80.03 %

We can see that from our plants convert coal to ethylene by five units has

mass efficiency as 45.47 % and total efficiency is 80.03%.

3.6. Energy Required per Unit Product and Energy EfficiencyFrom energy balance calculation above, we can calculate energy required

for a unit main product produced in this plant. It can be done by dividing amount

total energy required with main product produced.


86

Energy efficiency= Energy RequireEt h ylene Produced

=37,938,399,3361,378,076

Energy efficiency=27,080 kJ / Kg

We can see that from our plants convert coal to ethylene by five units has

energy efficiency as 27.080 kJ/kg.


CHAPTER IVCONCLUSION

1. Ethylene production is still limited compared to the increasing demand, while

coal as one of ethylene feedstocks is abundant in Indonesia.

2. Raw material for coal will be supplied from PT Bukit Asam and PT Adaro,

while oxygen will be supplied from PT Air Product’s integrated facility

within the plant.

3. The plant will be located in Cilegon.

4. The plant will consists of five units, which are Gasifier Unit, Acid Gas

Removal Unit, Water Gas Shift Unit, Methanol Synthesis Unit and Ethylene

Synthesis Unit.

5. Gasifier technology that is chosen in Entrained Bed Gasifier, while the

ethylene synthesis will be done using Methanol-to-Olefin technology.

6. The plant has mass efficiency of 45,74% while the energy efficiency of the

plant is 27.080 kJ/kg.


REFERENCES

Anonymous. 2014. www.maps.google.com. Accessed on 12 September 2014 at

20.10

Anonim. (2013). Ethylene via Ethanol Dehydration. Accessed from

http://base.intratec.us/home/chemical-processes/ethylene/ethylene-via-

ethanol-dehydration at 18th September 2014

Anonim. (2008). Methanol Plant Process Description. Accessed from

http://www.ippe.com/plants/600304/process_description.pdf at 12th

September 2014

Anonymous. 2014. Statistik Batubara. www.esdm.go.id. Jakarta: Direktorat

Jenderal Mineral dan Batubara Kementerian ESDM. Accessed on 12

September 2014 at 19.17

Asma. A. et al. (2008). Ethanol Conversion from Syngas-Microbial Conversion.

Accessed from

http://bioweb.sungrant.org/Technical/Biofuels/Technologies/Biofuels+from

+Syngas/Ethanol+from+Syngas/Default.htm at 18th September 2014

Fisher Controls International LLC. (2010). Chapter 1: Ethylene Production.

Accessed from www.fisher.com at 18th Sepetember 2014

Habiburrohman. Muhammad. 2012. Perancangan Pabrik Gas Sintesis

Menggunakan Proses Gasifikasi Batubara Sebagai Bahan Baku Pembuatan

Bahan Bakar Cair. Depok: Universitas Indonesia.

He. J. et al. (2008). Research on ethanol synthesis from syngas. Journal of

Zhenjiang University Science A. 2008 9(5):714-719

Sunggyu. L.. Speight. J. G.. Loyalka. S. K. (2007). Handbook of Alternative Fuel

Technology. Boca Raton: CRC Press


http://www.fisher.com/

http://bioweb.sungrant.org/Technical/Biofuels/Technologies/Biofuels+from+Syngas/Ethanol+from+Syngas/Default.htm

http://bioweb.sungrant.org/Technical/Biofuels/Technologies/Biofuels+from+Syngas/Ethanol+from+Syngas/Default.htm

http://www.esdm.go.id/

http://www.ippe.com/plants/600304/process_description.pdf

http://base.intratec.us/home/chemical-processes/ethylene/ethylene-via-ethanol-dehydration

http://base.intratec.us/home/chemical-processes/ethylene/ethylene-via-ethanol-dehydration

http://www.maps.google.com/

revisi tk11 report assignment 1

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