Production of 50 000 MT Per Year Butadiene

1

CHAPTER 1

INTRODUCTION: PROJECT CONCEPTION AND LITERATURE REVIEW

Butadiene is a versatile raw material used in the production of a wide variety of synthetic rubbers and polymer resins as well as a few chemical intermediates. The largest uses for butadiene are the production of styrene butadiene rubber (SBR) and polybutadiene rubber (BR), which are used mainly in tire products.[ Anonymous, (February 2009),Butadiene Uses and Market Data ]

Butadiene is one of the components used in the manufacture of acrylonitrile-butadiene-styrene (ABS), styrene-butadiene (SB) copolymer latex, styrene-butadiene block copolymers and nitrile rubbers. 1, 3-Butadiene ranks 36th in the most produced chemicals in the United States. Three billion pounds per year are produced in the United States and 12 billion globally. World butadiene consumption in the synthetic rubber and latex applications is forecast to grow at an average rate of about 2%/year.[ Anonymous, (February 2009),Butadiene Uses and Market Data]

The region seeing the strongest performance has been Asia due to increased production of finished goods in the electronics, automobile and tire

2 sectors. The major source of butadiene is as a byproduct in the steam cracking of naphtha and gas oil to make ethylene and propylene. The butadiene is extracted from the C4 cracker stream using extractive distillation. Butadiene is a colorless, non corrosive liquefied gas with a mild aromatic or gasoline-like odor. Butadiene is both explosive and flammable because of its low flash point.[ Anonymous, (February 2009),Butadiene CAS No: 106-99-0]

1.1)

History and Background

1.1.1) History

 In 1863, a French chemist isolated a previously unknown hydrocarbon from the pyrolysis of amyl alcohol. This hydrocarbon was identified as butadiene in 1886, after Henry Edward Armstrong isolated it from among the pyrolysis products of petroleum. In 1910, the Russian chemist Sergei Lebedev polymerized butadiene, and obtained a material with rubber-like properties. This polymer was, however, too soft to replace natural rubber in many roles, especially automobile tires.[ Anonymous, (February 2009),History Butadiene]

 The butadiene industry originated in the years leading up to World War II. Many of the belligerent nations realized that in the event of war, they could be cut off from rubber plantations controlled by the British Empire, and sought to remove their dependence on natural rubber. In 1929, Eduard Tschunker and Walter Bock, working for IG Farben in Germany, made a copolymer of styrene and butadiene that could be used in automobile tires. Worldwide

3 production quickly ensued, with butadiene being produced from grain alcohol in the Soviet Union and the United States and from coal-derived acetylene in Germany.[ Armstrong, H.E. Miller, A.K. (1886).]

1.1.2) Background

 1, 3-Butadiene is a simple conjugated diene. It is an important industrial chemical used as a monomer in the production of synthetic rubber. When the word butadiene is used, most of the time it refers to 1,3-butadiene.[ Sun, H.P. Wristers, J.P. (1992).]

 The name butadiene can also refer to the isomer, 1,2-butadiene, which is a cumulated diene. However, this allene is difficult to prepare and has no industrial significance.

 In the United States, western Europe, and Japan, butadiene is produced as a byproduct of the steam cracking process used to produce ethylene and other olefins. When mixed with steam and briefly heated to very high temperatures (often over 900 °C), aliphatic hydrocarbons give up hydrogen to produce a complex mixture of unsaturated hydrocarbons, including butadiene. The quantity of butadiene produced depends on the hydrocarbons used as feed. Light feeds, such as ethane, give primarily ethylene when cracked, but heavier favor the formation of heavier olefins, butadiene, and aromatic hydrocarbons.

4  Butadiene is typically isolated from the other four-carbon hydrocarbons produced in steam cracking by extraction into a polar aprotic solvent such as acetonitrile or dimethylformamide, from which it is then stripped by distillation.

 Butadiene can also be produced by the catalytic dehydrogenation of normal butane. The first such commercial plant, producing 65,000 tons per year of butadiene, began operations in 1957 in Houston, Texas.

 In other parts of the world, including eastern Europe, China, and India, butadiene is also produced from ethanol. While not competitive with steam cracking for producing large volumes of butadiene, lower capital costs make production from ethanol a viable option for smaller-capacity plants. Two processes are in use.

 In the single-step process developed by Sergei Lebedev, ethanol is converted to butadiene, hydrogen, and water at 400–450 °C over any of a variety of metal oxide catalysts:

 This process was the basis for the Soviet Union's synthetic rubber industry during and after World War II, and it remains in limited use in Russia and other parts of Eastern Europe. In the other, twostep process, developed by the Russian chemist Ivan Ostromislensky, ethanol is oxidized to acetaldehyde, which reacts with additional ethanol over a tantalum-promoted porous silica

5 catalyst at 325–350 °C to yield butadiene:[ Beychok, M.R. and Brack, W.J, June 1957]

 2 CH3CH2OH →  CH2=CH-CH=CH2 + 2 H2O + H2

Figure 1.1: Structural Chemical Reaction of Ethanol

 CH3CH2OH + CH3CHO →  CH2=CH-CH=CH2 + 2 H2O

Figure 1.2: Structural Chemical Reaction of Ethanol by react With tantalum-promoted porous silica

6  This process was used in the United States to produce government rubber during World War II, and remains in use today in China and India.

1.1.3) Butadiene Synonyms and Abbreviations  Biethylene  Buta-1,3-diene  Butadieno  Divinyl  Erythrene  Vinylethylene

1,3-Butadiene

1.1.4) Chemical-Physical Properties product and raw material 

Molecular formula C4H6



Molar mass 54.09 g mol−1



Appearance Colorless gas or refrigerated liquid



Density 0.64 g/cm at -6 °C, liquid



Melting point -108.9 °C, 164 K, -164 °F



Boiling point -4.4 °C, 269 K, 24 °F



Solubility in water 735 ppm



Viscosity 0.25 cP at 0 °C

7 1.1.5) Importance of Butadiene production

The 1,3-butadiene is the simplest member of the series of conjugated   dienes,   which   contain   the   structure   C=C−C=C,   the   C   being   carbon. The wide variety of chemical reactions peculiar to this system makes butadiene important in chemical synthesis. Under the influence of catalysts, butadiene molecules combine with each other or with other reactive molecules, as acrylonitrile or styrene, to form elastic, rubberlike materials. In uncatalyzed reactions with reactive unsaturated compounds, such as maleic anhydride, butadiene undergoes the Diels-Alder reaction, forming cyclohexene derivatives. Butadiene is attacked by the numerous substances that react with ordinary olefins, but the reactions often involve both double bonds (e.g., addition of chlorine yields both 3,4-dichloro-1butene and 1,4-dichloro-2-butene). At atmospheric conditions, 1,3butadiene exists as a colourless gas, but it is liquefied either by cooling to -4.4° C (24.1° F) or by compressing to 2.8 atmospheres at 25°C. [Kirshenbaum, I. (1978)]

1.2)

Application of Products

Nearly all (96%) of the butadiene produced globally is as a co-product of the steam cracking of naphtha and gas oil to make ethylene and propylene. After ethylene  and  propylene  are  extracted  from  the  cracker,  a  “C4  stream”  is  separated   from the process which contains predominately hydrocarbons containing four carbon atoms, e.g. butadiene and butenes.

The largest single use for butadiene is in the production of styrenebutadiene rubber (SBR) which, in turn, is principally used in the manufacture of

8 automobile tyres. SBR is also used in adhesives, sealants, coatings and in rubber articles like shoe soles. Polybutadiene is also used in tyres and can be used as an intermediate in the production of acrylonitrile-butadiene-styrene (ABS). ABS is widely used in items such as telephones, computer casings and other appliances.[ Anonymous, (June 21, 2007),Product Safety Assessment, Butadiene]

Other polymers made from butadiene include styrene-butadiene latex, used for example in carpet backings and adhesives; nitrile rubber, used in hoses, fuel lines, gasket seals, gloves and footwear; and styrene-butadiene block copolymers which are used in many end-uses ranging from asphalt modifiers (road and roofing construction applications), to adhesives, footwear and toys.[ Anonymous, (June 21, 2007),Product Safety Assessment, Butadiene]]

Chemical intermediates made from butadiene include adiponitrile and chloroprene which are used, respectively, in the manufacture of nylon and neoprene.

Figure 1.3: Chart of uses 1,3 Butadiene

9 1.2.1) Synthetic Elastomer

The synthetic elastomers of the invention have incorporated therein from about 11-50%, preferably from about 20-40%, of a liquid, high vinyl 1,2polybutadiene resin having a pendant vinyl group for every other chain carbon which is capable of crosslinking to a very high degree. The preferred liquid, high vinyl 1,2 polybutadiene has from about 80-95 mole %, most preferably from about 90-95 mole % 1,2 vinyl structure. [Anonymous, (1987), Synthetic elastomeric with improved chemical, aging and oil resistance]

In the method of the invention, the previously polymerized, liquid, high vinyl content 1,2-polybutadiene is incorporated into an elastomer selected from the group consisting of ethylene-propylene copolymer rubbers and ethylenepropylene-nonconjugated diene terpolymer rubbers. The previously polymerized liquid, high vinyl content 1,2-polybutadiene is incorporated during the polymerization of the elastomer to provide additional cure sites on the resulting elastomer. Rather than attempting to directly polymerize the polybutadiene onto the backbone of the ethylene-propylene chain, the polybutadiene is solution blended after catalysis and prior to separating and drying the polymerized elastomer. The polybutadiene is added prior to precipitating and drying the polymerized elastomer. The resulting elastomer is peroxide cured to produce an insulating material exhibiting excellent electrical characteristics, ease of compounding, and improved performance at extreme temperatures and pressures when exposed to solvents, oil and aqueous environments.[ Anonymous, (1987), Synthetic elastomeric with improved chemical, aging and oil resistance]

10 Synthetic olefin polymers are popular as electrical insulating materials because of their ease of compounding, good extrudability and excellent electrical characteristics. These polymers also find use as valve seats, and in other applications. In particular, ethylene-propylene copolymer rubbers, known as EPR, and ethylene-propylene-nonconjugated diene terpolymer rubbers, known as EPDM have been widely employed as the primary insulating materials for electrical wire and cable. These materials have the characteristics of flowing and/or distorting at elevated temperatures and under extreme pressures and are sensitive to swelling and dissolving in various hydrocarbon solvents and oils. Where insulated wire and cable is needed for extreme conditions, EPR and EPDM elastomers have been physically blended with low molecular weight polybutadiene in a roll mill, Banbury mixer, or the like. The physical blending or incorporation of the polybutadiene into the EPR/EPDM rubber provides additional cure sites for greater cross link density. An increase in cross link density has been found to enhance the chemical aging and oil resistance of the elastomer, improving the performance of the elastomer in extreme environmental conditions. U.S. Pat. No. 3,926,900 to Guzy et al., issued Dec. 16, 1975, discusses the physical blending of liquid 1,2 polybutadiene with EPDM polymers. [Anonymous, (1987), Synthetic elastomeric with improved chemical, aging and oil resistance]

1.2.2) Polymer and Resin

Engineering resins is the term for a group of polymer plastics which exhibit a greater tendency to form crystals in their solid state than their more amorphous cousins. The additional level of long-range order at the molecular

11 scale produces a different set of physical properties which suit the engineering plastic resins to a wide variety of applications that amorphous resins cannot fill. In general, engineering plastic resins are physically stronger and less flexible than amorphous resins and show greater resistance to fatigue, friction and wear. [Anonymous, (2007), Engineering Resin]

1.2.3) Polybutylene Terephthalate (PBT)

PBT engineering plastic resins are used to fabricate components found in computer keyboards, appliances, fluid handling systems, cars and trucks, electrical connectors, and industrial systems and controls. This product list is a testament to the versatility of the compound and is a direct result of its many outstanding characteristics. Stability and resistance to temperature extremes, along with a superior ability to be molded into complex or fine shapes makes PBT one of the most important engineering polymers.[ Anonymous, (2007), Engineering Resin]

1.2.4) PC/ABS A true industrial thermoplastic, this engineering resins blend combines the most desirable properties of both materials; excellent features of ABS and the superior mechanical properties and heat resistance of polycarbonate. PC-ABS blends are widely used in automotive, electronics and telecommunications applications. This engineering plastic resins blend is ideal for the rapid production

12 of prototypes, tooling and the direct (tool-less) manufacturing of production parts.[ Anonymous, (2007), Engineering Resin]

1.2.5) Nylon 66 (Polyamide 66) resin

A thermoplastic resin with excellent mechanical, thermal and electrical properties will use as raw materials of fiber, film and engineering plastic. Engineering plastic resins are replacing the previous metals at a rapid pace. Nylon has a proven record of outstanding service in a wide range of applications for all industries.[ Anonymous, (2007), Engineering Resin]

1.2.6) Styrene Butadiene Rubber (SBR)

Styrene butadiene rubber (SBR) is the outcome of synthetic rubber research that took place in the United States and Europe under the impact of the shortage of natural rubber, a German chemist developed a series of synthetic elastomers by copolymerization of two compounds (styrene and butadiene) in the presence of a catalyst. The first step involved in the process is to let styrene and butadiene react together. The new synthetic rubber that was formed consists of about 25% styrene, with butadiene making up the rest, which in principle had the same properties as natural rubber. This rubber is considered to be the highest volume general purpose and the most common type of synthetic rubber. [Anonymous, (2007), Types of Synthetic Rubber]

13 1.2.7) Properties of Styrene Butadiene Rubber



This type of rubber is usually very weak unless reinforcing fillers are incorporated. With suitable fillers, this becomes a strong rubber.



It has similar chemical and physical properties like natural rubber.



It has better abrasion resistance.



It has poorer fatigue resistance.



Heat resistance is better than natural rubber.



Low temperature flexibility and tensile strength are less than that of natural rubber.

1.2.8) Chemical used

Chemical intermediates manufactured from butadiene include adiponitrile and chloroprene. Adiponitrile is used to make nylon fibres and polymers. Chloroprene is the monomer to make polychloroprene, better known as Neoprene, which has a wide variety of uses such as wet suits, electrical insulation, car fan belts, gaskets, hoses, corrosion-resistant coatings and adhesives.[ Anonymous, (February 2009),Butadiene Uses and Market Data]

1.2.9) Other applications Elastomers, 61% (styrene-butadiene rubber (SBR), 32%; polybutadiene, 23%; polychloroprene, 4%; nitrile, 2 percent); styrene-butadiene latex, 12%; adiponitrile for HMDA, 11%; ABS resins, 5 percent; miscellaneous, 11%

14 Anonymous, (November 1996),Locating and Estimating Air Emissons From Source]

Other polymers made from butadiene include styrene-butadiene (SB) copolymer latex, which is used in paper coatings, carpet back coatings, foam mattresses and adhesives. Styrene-butadiene block copolymers have many applications ranging from asphalt modifiers in road and roofing construction to adhesives, footwear and toys.

Nitrile rubbers, made by the copolymerisation of acrylonitrile with butadiene, are used mainly in the manufacture of hoses, gasket seals and fuel lines for the automobile industry as well as in gloves and footwear.

1.3.

Problem Statement Butadiene is one of highly demanded products in petrochemical industry.

For many years, its production rate has been increasing. The current production of butadiene is about 7,000,000 ton per year in USA, Western Europe and Eastern Asia only, and it does not satisfy the market needs, since yearly increase in demand is predicted to be 3.9%, whereas increase of production rate is 2-3% only. The price of the product during 2000 increased by 25%. Butadiene is produced using n-butane as a raw material in a two stage Gudry vacuum dehydrogenation process. The output of butadiene in this process is usually about 12%. The project presents the extremely effective solution for production butadiene – the catalyst that makes it possible to increase output of butadiene from 12 to 25%.[DR. Talishinsky, (1996),Butadiene production]

15 Besides elastomers will continue to be the largest consumer of butadiene and should maintain their position of 61 percent of total consumption. However, they are mature products that are heavily reliant on the automotive industry. Adiponitrile/ hexamethylenediamine (HMDA), styrene-butadiene (SB) copolymer latex, acrylonitrile-butadiene-styrene (ABS) resins, styrenic block copolymers and other smaller polymer applications will grow faster than the elastomers (excluding polybutadiene), but they each account for only 5-10% of the total butadiene market. With a projected negative average annual growth of -1.7 during 20002004, the total market for butadiene in 2004 will reach 5.1 billion pounds, or about about what it was in 1998. This takes into account the big hit in demand in 2001.

[Lynne

M.Miller,

(Dec

1978),Investigate

of

Selected

Potential

Environmental Contaminants : Butadiene and its Oligomers,]

So to recover the quality and maintain the production cause of the high of demand in Malaysia and the entire world, the selected of this plant design research title are very suitable.

1.4.

Objective and Scope

The objective of this research of plant design is to increase the production of butadiene with efficient way and to bear an amount of demanding production especially in Malaysia with the scope of this research are:-

i.

To design the plan based on demand of production

ii.

To develop a suitable business in Malaysia

iii.

To make a profit from the production

16 1.5)

MARKET SURVEY

1.5.1) Global Situation

The Global production and consumption of butadiene in 2008 were approximately 10.6 million metric tons and 11.1 million metric tons, respectively. Global capacity utilization in 2008 was 88%. Global butadiene consumption is estimated to have increased by almost 2% in 2008, and is expected to average growth of 3.8% per year from 2008 to 2013, slowing to 2.3% per year from 2013 to 2018. Global utilization rates are expected to be in the 90s. [Anonymous, (January 2010),Butadiene]

Styrene butadiene rubber (solid & latex) accounted for more than 30% of global butadiene consumption in 2008, followed by polybutadiene rubber, for around 25%. Other applications for butadiene include manufacture of styrenic copolymers, ABS resins, SB latex, nitrile rubber, and adiponitrile/HMDA. The following pie chart shows world consumption of butadiene by end use: [Anonymous, (January 2010),Butadiene]

From  the   figure  1.4,   it’s  shown  that  butadiene  demand   is  concentrated  in   its use in the manufacture of styrene butadiene rubber (SBR) solid and latex (34.7%), polybutadiene rubber (24.9%), ABS resins (10.2%), SB copolymer latex (9.4%) and other consumptions about 20.9%.

17

Figure 1.4: World consumption of Butadiene

Table 1.1.0: World Butadiene Supply/Demand Balance (1999-2005)

Demand for butadiene in the production of ABS resins will see the highest average annual rate growth for all derivatives in the increase in total tons of butadiene consumed. Demand of butadiene in this application will increase by more than 500,000 tons during the period. Global demand for butadiene will increase at an average annual rate of 3.9% during the period from 2001-2006 percent and will outpace capacity additions. This rate is higher than the

18 compounded annual rate of 2.7% from 1996-2001 due to the global decline in demand that occurred in 2001 following the global economic slowdown.[ Jorg Wutke, (1996),The petrochemical Industry in China]

It’s   expected   that,   in   2008   through   2012   period   will   experience   a   butadiene demand growth rate of just under 3.5 percent per year, slightly higher than the 3.2 percent annual rate experienced over the past five years. Global demand for butadiene consumed into ABS resin production is estimated to grow at a high annual rate of around five percent, due to heavy use of thermoplastics in the manufacture of computer equipment and other appliances, mainly in China. Butadiene based nylon production, through adiponitrile, will also grow at about five percent per year. However, worldwide demand for butadiene in its largest end use sector, the production of commodity-based synthetic rubber and latex, is anticipated to average around 3% per year.[ Anonymous, (January,14,2008), CMAI Completes 2008 World Butadiene Analysis]

1.5.1.1)Styrene Butadiene Rubber (SBR) Demand

The tyre industry consumes 75 percent of the SBR produced globally followed by the mechanical rubber goods/automotive parts applications (19 percent of the market). Footwear accounts for only around six percent of the SBR market. The main use of SBR is in the manufacture of tyre tread, and consumption is forecast to develop in line with the automotive sector.

1. The production of auto tyres is increasingly competitive and cost sensitive. Consequently, the manufacture of tyres and other rubber goods has tended

19 to migrate to lower labour cost areas, depressing market growth in developed regions such as Western Europe, the United States and Japan. Exports of finished rubber goods, primarily tyres, from regions such as China to the United States and Europe have increased dramatically over the last five years, leading to the closure of a number of tyre plants in the importing regions. Flourishing automotive sectors in China, India, and Thailand have also increased demand in the Asia Pacific region. [Anonymous, (2008), Butadiene Derivatives Impacted by Automotive Crisis]

Figure 1.5: Global SBR Capacity Additions/ (Closures)

20 1.5.1.2 Butadiene Rubber (BR)

Approximately two-thirds of BR is consumed in tyre production, with a further quarter used as an impact modifier in high impact polystyrene (HIPS) production. Other applications consume only around eight percent of the BR market. As the main use of BR is in the manufacture of tyres, BR consumption is forecast to increase in line with the automotive sector. Asia Pacific, North America and Western Europe are the major consuming regions for BR, with total consumption in these three regions accounting for more than 80 percent of the global total. China has surpassed the United States to become the largest consumer of BR in 2007. The combination of new tyre manufacturing and high impact polystyrene (HIPS) capacities in China has boosted demand for BR while some rationalisation of both capacities was seen in the United States.[ Anonymous, (1996), Butadiene rubber]

Figure 1.6: Global Butadiene Rubber Capacity Additions/ (Closures)

21 1.5.1.3 Acrylonitrile Butadiene Styrene (ABS)

Global ABS demand has been under pressure from inter-polymer competition, especially from polypropylene and lately polystyrene, which is competing particularly at the lower specification end of the automotive sector. Recent development in high gloss polystyrene is a new threat for ABS for decorative parts. However, ABS remains the material of choice in most applications in the key electronics/electrical appliance sector, due to its mechanical properties, high gloss and processability.[ Anonymous, (November 1996),Butadiene Styrene]

Although ABS consumption is forecast to grow at slower rates over 20092018 after the recent economic downturn, it will be one of the key drivers for styrene market growth during the recovery of the economy, with long term sustainable growth supported by the electrical appliance and automotive sectors. Asia Pacific, particularly China will remain the largest consuming region with an increasing proportion of the global consumption. Central Europe is expected to grow to balance the slowdown in the Western Europe.

22

Figure 1.7: Global ABS Capacity

In 2008, SBR is the largest end use of butadiene, accounting for slightly less than one-third of total demand, followed by BR and SBL respectively. ABS, hexamethylenediamine (adipic acid) HMDA and other butadiene uses made up the remaining demand, accounting for 30 percent in total. Butadiene consumption is driven to a great extent by the automotive industry, which tends to give a very volatile growth pattern. Historically, BR grew faster than SBR, but this will change in the forecast due principally to the slow growth in HIPS market. Despite a freefall in the ABS sector last year along with electronics and automotive industries, ABS is expected to recover and continues to grow at high rates. Due to a cost advantage over the acrylonitrile process, demand into HMDA towards butadiene will also grow rapidly as new plants start up in the United States and China. On the other hand, the growth of the SBL sector is forecast to moderate as a result of more efficient use in paper and carpet industries. [Anonymous, (2007),Product overview and market projection of emerging bio]

23

Figure 1.8: Global Butadiene Capacity

Butadiene extraction capacity is concentrated in the major naphtha cracking regions of Asia Pacific, North America and Western Europe. The development of ethylene capacity based on heavier feedstocks in the Middle East will increase butadiene capacity there, although the region is destined to remain small in terms of overall production. Capacity in Eastern Europe is expected to remain fairly flat as ethylene capacity in the region remains in excess of demand, and is not expected to increase significantly.[ Anonymous, (2009),Butadiene Market Dynamics]

24

Figure 1.9: Global Butadiene Consumption, Operating Rate and Capacity

Global butadiene operating rates remain at well above average levels, but are expected to decline towards a trough in 2011 as major capacity additions are made during the forecast period of low demand growth. The increasing proportion of liquids based cracker developments will increase the availability of mixed C

4

feedstock for butadiene extraction at a rate greater than that of

butadiene demand growth. This is expected to result in a greater proportion of mixed C 4 hydrogenation and co-cracking rather than over expansion of butadiene extraction capacity. The limited amount of naphtha cracker capacity expansion in North America and Western Europe will govern the level of butadiene capacity development in these areas. No new derivatives will be based in areas where there is no additional butadiene availability, leading to a gradual concentration of activity in butadiene and derivatives in Asia. [Anonymous, (2009),Butadiene Market Dynamics]

25 Year

Price ($/Pound)

2004

0.30

2005

0.26

2006

0.544

2007

0.735

2008

1.360

2009

0.428

Table 1.1.1: World Butadiene Prices (2004-2009)

Figure 1.10: Global Butadiene Prices (2004-2009)

26 Year

Demand (-000-Metric Tons)

1999

7,880

2000

8,340

2001

8,634

2002

8,937

2003

9,229

2004

9,507

2005

9,810

2006

10,430

2007

10,878

2008

11,513

Table 1.1.2: World Demand toward Butadiene

Figure 1.11: World Demand toward Butadiene (1998-2010)

27 According to the figure 1.4.5, world demand towards butadiene was slowly increased from 1999 to 2008. In 2009, global butadiene demand is expected to grow at a pace lower than the 3.2 percent annual rate experienced over the past five years. For example, the outlook for worldwide butadiene in its largest end use sector, the production of commodity based synthetic rubber and latex, is anticipated to only average around 2 percent per year. A slowing global economy is also causing slower demand for rubber goods, especially in the automotive sector. Global butadiene growth has averaged 3.3 percent per year from 1995-2006, but is expected to average only 3.1 percent over 2006-2015. Global consumption of butadiene is expected to increase from 10 million tons in 2006 to 13 million tons by 2015.[ Anonymous, (2008),Basic Material: Global Insights]

From figure 1.4.4, the global prices of butadiene were rapidly increasing from 2004 to 2008 but the price was drastically decreased in 2009. Actually, the global economic culture and oil prices were affecting the prices of butadiene in the market. We believe that the global prices rhythm of butadiene will increased according to the report that said the global economic will became stable at the end of 2010. The political instability especially in the Middle East (Iraq and Iran) will cause the increasing of global oil prices. So, we assumed that when the oil prices increase, the global prices of butadiene will increased too.

1.5.2 Asia Pacific Situation

The Asian market has been particularly active in building new capacity of butadiene and butadiene derivatives due to the ongoing development of automotive and tyre production in the region. The relocation of automotive industries increased synthetic rubber demand through tyre production, while both ABS and HMDA will benefit from plastics demand in the Asian automotive

28 sector. Additional global demand for butadiene in recent years was entirely focussed in Asia Pacific where significant new derivatives capacity built up, particularly in China and South Korea. In the outlook, the share of Asia Pacific demand will grow further from 45 percent in 2008 to 53 percent in 2015.

Demand in Asia Pacific accounts for 41 percent of the global total, and the proportion is forecast to increase. The growth in demand in Asia is driven by increasing availability, and the rapid growth in demand in derivatives to supply the booming Asian manufacturing sector.[ Anonymous, (2008),Basic Material: Global Insights]

The markets for butadiene have emerged from a long period of oversupply, leading to record prices and margins in 2006. Butadiene prices broke the $1,500 per ton level in Asia in late 2006, almost six times the lowest prices seen in the late 1990s. At the same time, margins for West European producers reached over $400 per ton, despite the prevailing high feedstock prices. The current high global operating rates are set to last through 2007 and 2008, before dropping off due to major capacity additions.

The Asian market has been particularly active for butadiene due to the ongoing development of automotive and tyre production in the region which drives demand, and the major steam cracker developments which drive supply. The Asia Pacific region has accounted for over half of global capacity and demand growth over the past five years, and will account for three quarters over 2007-2011.

29 Major steam cracker developments in Asia will provide more mixed C4s for butadiene production, leading to a decrease in operating rates. In the long term, the growth of ethylene production, and therefore the availability of C4s, will exceed demand growth for butadiene, leading to increased reprocessing of steam cracker C4s. Countries such as China are expected to extract enough butadiene to serve their own derivative requirements, and hydrogenate then recycle the remainder back to the steam cracker.[ Anonymous, (2008),Basic Material: Global Insights]

The operating rates will remain above 85% until 2009, when major capacity additions and slower demand growth will cause a decline towards a trough in 2011. New capacity developments are focussed on conventional extraction from steam cracker mixed C4 streams. The current high margins on butane dehydrogenation are expected to be temporary and not likely to encourage new investment in this technology. The tendency towards heavier cracker slates in the Middle East is increasing the availability of steam cracker C4s for butadiene and derivatives.

The scale however remains small relative to

expansions in Asia . While currently growing rapidly from a small base, the Middle East is unlikely to build a major export position for butadiene and derivatives as it has in the ethylene chain.[ Anonymous, (2008),Basic Material: Global Insights]

1.5.2.1 Butadiene Market in China With the increase of domestic butadiene production capacity, China's butadiene supply will basically meet the rising demand in the coming four years, according to industry experts.With large scale development of the ethylene industry, enterprises under the aegis of China's two oil giants PetroChina and Sinopec are swarming to build or expand butadiene production facilities to

30 produce butadiene, which is in great demand on the domestic market. It is predicted   that   China’s   butadiene   production   capacity   is   expected   to   reach   2.7   million tons by 2011. [Anonymous, (2008),Management Discussions]

China had 18 butadiene producers and 26 sets of butadiene production facilities with an annual production capacity totaling 1.614 million tons by May 2007, accounting for 13.5 per cent of the world's total. Last year, it produced 1.153 million tons of butadiene, an increase of 15.78 per cent over 2005. Its butadiene output grew at an average annual rate of 12.3 per cent in 2001-2006. However, its current output cannot satisfy the domestic demand. It has to import butadiene in bulk. With some production facilities newly built or expanded, the import volume has declined moderately, from 195,900 tons in 2004 to 147,200 tons in 2005 and 89,200 tons in 2006. With the rapid development of synthetic rubber industry, the main consumer of butadiene, China's apparent butadiene consumption has kept growing in recent years, from 782,500 tons in 2001 to 1.0353 million tons in 2004 and 1.2153 million tons in 2006. It is predicted that the consumption will grow 8.7 per cent annually from 2006 to 2011, topping 1.7 million tons in 2011.[ Anonymous, (Nov 2008),Production from China]

1.5.3 Malaysia Situation Malaysia has a well-developed oil and gas sector and a growing petrochemical industry. The petrochemical industry is an important sector in Malaysia with investments totaling US$7.4 billion in 2004 and US$6.9 billion in 2007. From being an importer of petrochemicals, Malaysia is today an exporter of major petrochemicals product. A wide range of petrochemicals are produced in Malaysia such as olefins, polyolefin, aromatics, ethylene oxides, glycols, oxo-

31 alcohols, exthoxylates, acrylic acid, phthalic anhydride, acetic acid, styrene monomer, polystyrene, ethyl benzene, vinyl chloride monomer and polyvinyl chloride.[ Anonymous, (2009),Butadiene Market Dynamics]

The rapid growth of the industry is mainly attributed to the availability of oil and gas as feedstock, a well-developed infrastructure, a strong base of supporting services, the country's cost competitiveness, as well as Malaysia's strategic location within ASEAN and its close proximity to major markets in the Asia Pacific Region. Malaysia has the world's 14th largest natural gas reserves and 23rd largest crude oil reserves. In 2008, Malaysia produced 5,891 million standard cubic feet per day of natural gas and 691,600 barrels of oil equivalent per day of crude oil. Malaysia also has the world's largest production facility at a single location of liquefied natural gas with production capacity of 23 million metric tonne per year.[ Anonymous, (2009),Butadiene Market Dynamics]

The long term reliability and security of gas supply ensures the sustainable development of the country's petrochemical industry. The existence of a transpeninsular gas transmission pipeline system and six gas processing plants, has resulted in a ready supply of gas to the industry. To complement the existing gas reserves and to ensure further security of gas supply, Malaysia has forged partnerships with other ASEAN members for the supply of gas such as Vietnam, Indonesia and the Malaysia-Thailand Joint Development Area (JDA). In addition, gas supply will be further enhanced with the implementation of the ASEAN gas grid, a venture to make gas available to all the 10 ASEAN countries. With the full implementation of AFTA, petrochemical manufacturers in Malaysia will benefit from a single market. Manufacturers based in Malaysia will also benefit from the access to a much larger Asia Pacific market. With China being a net importer of petrochemicals, Malaysia's 'early harvest' Free Trade Agreement with China will

32 open up new business opportunities for petrochemicals manufacturers in Malaysia.[ Anonymous, (2009),Butadiene Market Dynamics]

The presence of world renowned petrochemical companies, such as Dow Chemical, BP, Shell, BASF, Eastman Chemicals, Toray, Mitsubishi, Idemitsu, Polyplastics, Kaneka, Dairen and West Lake Chemical speaks clearly of Malaysia's potential as an investment location for petrochemical industries. Most of these companies are working in collaboration with Malaysia's national petroleum company, PETRONAS. Three major petrochemical zones have been established in Kertih, Terengganu; Gebeng, Pahang; and Pasir Gudang/Tanjung Langsat, Johor. Each zone is an integrated complex with crackers, syngas and aromatics facilities to produce feedstocks for downstream products. There are also other petrochemical plants in Malaysia such as the ammonia and urea plants in Bintulu, Sarawak and Gurun, Kedah; acrylonitrile butadiene styrene plant in Pulau Pinang; methanol plant in Labuan and the nitrile-butadiene rubber plant in Kluang, Johor.[ Anonymous, (2009),Butadiene Market Dynamics]

1.6)

SCREENING OF SYNTHESIS ROUTE IN PRODUCTION OF

BUTADIENE

Butadiene is produced commercially by three processes: 

Steam Cracking of Paraffinic Hydrocarbons



Catalytic Dehydrogenation of n-Butane and n-Butene (the Houdry process).



Oxidative Dehydrogenation of n-Butene (the Oxo-D or O-X-D process).

33 1.6.1) Butadiene Production via Steam Cracking of Paraffinic Hydrocarbons

Figure 1.12: Typical Olefin Plant

In this process, butadiene is a co product in the manufacture of ethylene (the ethylene co-product process).The steam cracking process is reported to be the predominant method of the three processes of production, accounting for greater than 91% of the world's butadiene supply. Figure 1.1 depicts a flow chart for a typical olefins plant. The flow path of the C4 components (including butadiene) is indicated by bold [red] lines.

The indicated feed stocks (ethane, propane, butane, naphtha and gas oil) are fed to a pyrolysis (steam cracking) furnace where they are combined with steam and heated to temperatures between approximately 1450-1525 °F (790-830 °C). Within this temperature range, the feedstock molecules "crack" to produce

34 hydrogen, ethylene, propylene, butadiene, benzene, toluene and other important olefins plant co-products. After the pyrolysis reaction is quenched, the rest of the plant separates the desired products into streams that meet the various product specifications. Process steps include distillation, compression, process gas drying, hydrogenation (of acetylenes), and heat transfer. The focus of this review is 1,3butadiene; however, since butadiene is created in the olefins plant pyrolysis furnace, and is present in the crude butadiene product stream at concentrations up to approximately 75 wt%, the olefins plant process and the crude butadiene stream are addressed in this publication to a limited degree.[ Anonymous, (2002),Butadiene product Stewardship Guidance Manual]

While some olefins plant designs will accommodate any of the listed feed stocks, many olefins plants process only Natural Gas Liquids (NGLs) such as ethane, propane and sometimes butane. The mixes of feed stocks, the conditions at which the feed stocks are cracked, and the physical plant design, ultimately determine the amount of each product produced, and for some of the streams, the chemical composition of the stream. Olefins plants generally produce crude butadiene streams that contain very few C3 and C5 components, as shown by the analysis found in Table 1.1. The composition of the crude butadiene stream also can be altered via recycle blending of various product streams. For example, when finished butadiene streams (99+ wt% pure) do not meet commercial specifications, they are often combined with crude butadiene streams in order to recover the butadiene. In this situation, the resulting stream may not fall into the example range. Generally, crude butadiene is stored as a liquid under pressure in a pressure products sphere.[Anonymous, (2002),Butadiene product Stewardship Guidance Manual]

35

Table1.1.3: Example of a Crude Butadiene Analysis

1.6.2) Butadiene Production via Catalytic Dehydrogenation of n-Butane and n-Butene (the Houdry process)

The catalytic dehydrogenation of n-butane is a two-step process; initially going from n-butane to n-butenes and then to butadiene. Both steps are endothermic. A major butane-based process is the Houdry Catadiene process outlined in Figure 1.13. In the Houdry process, n-butane is dehydrogenated over chromium/alumina catalysts. The reactors normally operate at 12-15 centimeters Hg absolute pressure and approximately 1100-1260 °F (600-680 °C). Three or more reactors can be used to simulate continuous operation: while the first reactor is on-line, the second is being regenerated, and the third is being purged prior to regeneration. Residence time for feed in the reactor is approximately 5-15 minutes. As the endothermic reaction proceeds, the temperature of the catalyst bed decreases and a small amount of coke is deposited. In the regeneration cycle, this coke is burned with preheated air, which can supply essentially all of the heat required to bring the reactor up to the desired reaction temperature.[Anonymous, (2002),Butadiene product Stewardship Guidance Manual]

36

Figure 1.13: Catadiene Process Plant

The reactor effluent goes directly to a quench tower, where it is cooled. This stream is compressed before feeding an absorber/stripper system, where a C4 concentrate is produced to be fed to a butadiene extraction system for the recovery of high purity butadiene.

1.6.3) Butadiene Production via Oxidative Dehydrogenation of n-Butenes (the Oxo-D or O-X-D process)

Oxidative dehydrogenation of n-butenes has replaced many older processes for commercial (on-purpose) production of butadiene. Several processes and

many catalyst

systems

have been developed

for

the

oxydehydrogenation of either n-butane or of n-butene feed stocks. Butenes are much more reactive, however, and they require less severe operating conditions than that of n-butane to produce an equivalent amount of product. Therefore, the use of n-butane as a feedstock in this process may not be practical. In general, in

37 an oxydehydrogenation process, a mixture of n-butenes, air and steam is passed over a catalyst bed generally at low pressure and approximately 930 1110 °F (500-600 °C).

The heat from the exothermic reaction can be removed by circulating molten heat transfer salt, or by using the stream externally for steam generation. An alternate method is to add steam to the feed to act as a heat sink. The heat can then be recovered from the reactor effluent. Reaction yields and selective can range from 70-90%, making it unnecessary to recover and recycle feedstock. (Yield losses can produce the CO2.) In the Oxo-D process shown in Figure 1.3, a mixture of air, steam, and n-butenes is passed over the dehydrogenation catalyst in a continuous process. The air feed rate is such that an oxygen/butene molar ratio of approximately 0.55 is maintained, and the oxygen is totally consumed. A steam to butene ratio of 10:1 has been reported as necessary to absorb the heat of reaction and to limit the temperature rise.[Anonymous, (2002),Butadiene product Stewardship Guidance Manual]

The reactor effluent is cooled and the C4 components are recovered in an Absorber/degasser/stripper column combination. The lean oil flows from the bottom of the stripper back to the absorber, with a small amount passing through a solvent purification area. Crude butadiene is stripped from the oil, recovered in the overhead of the stripper, and then it is sent to a purification system to recover the butadiene product.[Anonymous, (2002),Butadiene product Stewardship Guidance Manual]

38

Figure 1.14: Oxidative Dehydrogenation Process

1.6.4) Butadiene Recovery from Crude Butadiene Streams Via Extractive Distillation

Since the boiling points of the various C4 components are so close to each other, separation via simple distillation does not currently suffice to adequately separate the components; therefore, extractive distillation is used. Several design options are available, including those listed in Table 1.2. Inclusion here is not intended as an endorsement. These processes involve one or two extractive distillation steps followed by one or two distillation steps. The number of extraction and/or distillation steps can be reduced to one by including an acetylene

hydrogenation

step.[Anonymous,

Stewardship Guidance Manual]

(2002),Butadiene

product

39

Table 1.1.4: Major Butadiene Recovery Process

1.6.5) Butadiene Purification via Acetylene Hydrogenation and Extractive Distillation Using MOPN/Furfural Solvent

This process contains four sections: 1) acetylene hydrogenation, 2) extractive distillation, 3) butadiene purification, and 4) solvent purification.

The objective of the acetylene hydrogenation section is to hydrogenate C4 acetylenes that could otherwise contaminate the butadiene product. This is achieved using a liquid phase reactor system. Butadiene-dimers and trimers formed in the reactor are removed via distillation in the green oil column located just downstream of the reactor. The green oil column overhead stream is fed to the extractive distillation section. The function of the extractive distillation section is to separate the C4 hydrocarbon stream into a butane/isobutene/transbutene-2 stream (C4 Raffinate 1) and a butadiene/cis-butene-2 stream via extractive distillation and solvent stripping. The green oil column overhead stream is vaporized then fed to the lower portion of the extraction column where the vapors are counter currently contacted with the aqueous methoxy-proprionitrile (MOPN)/furfural solvent which are fed into the top of the column. Butane and the less soluble butenes are concentrated and removed in the overhead stream.

40 The butadiene/cis-butene-2 rich solvent from the bottom of the extraction column are fed to the extract stripper column, where butadiene, cis-butene-2 and acetylenes (ppm level) are stripped overhead. The extract stripper column overhead stream is used to feed the butadiene purification column where butadiene is concentrated in the overhead product. Then the remaining butene-2 and heavier components are drawn from the bottom of the column and recycled to the olefins plant cracking furnaces. The purpose of the solvent purification section is to remove impurities from the lean solvent. The system consists of two evaporators, a stripping column and a solvent settling drum which are used to remove

furfural-butadiene

vinylcyclohexene

polymer,

compounds.

acrylonitrile-butadiene

[Anonymous,

codimer,

(2002),Butadiene

and

product

Stewardship Guidance Manual]

Figure 1.15: Process A- Acetylene Hydrogenation/Extractive Distillation Using MOPN/Furfural

41 1.6.6) Extractive and Conventional Distillation Process Using NMP Solvent

This process, licensed by BASF and illustrated in Figure 1.6, is a combination of extractive and conventional distillation. The extractive distillation uses n-methylpyrrolidone (NMP) as the solvent. The highest temperature is approximately 300 °F and the maximum pressure is approximately 100 psig (7 bars g).

The evaporated C4 cut is fed to the extractive distillation section where in the first stage the butanes and the butenes are separated from the more soluble 1,3-butadiene, 1,2-butadiene, C4 acetylenes, propyne and the C5 hydrocarbons. The loaded solvent is degassed in a steam heated column where the acetylenes are withdrawn as a side stream and are fed to a washer where the NMP is recovered.

Crude butadiene leaves the extractive distillation section at the top of the second stage and is then fed to the propyne distillation column where propyne (methyl-acetylene) is removed overhead. The bottoms product containing the 1,3butadiene, 1,2-butadiene and the C5 hydrocarbons is then distilled in the butadiene column. Generally, 1,3-butadiene with a purity of >99.6% by weight leaves the top of the butadiene column. The column bottoms stream usually contains 1,2-butadiene and heavier hydrocarbons.

The crude butadiene, the top of the propyne column, and the purified butadiene are typically inhibited with tertiary butyl catechol (TBC) or with other compounds. Sodium nitrite can be used as an inhibitor during extractive distillation. The waste hydrocarbon streams can be diluted with naphtha and used as supplemental feedstock for the olefins plant.[ Anonymous, (November 1996),Produce of Butadiene]

42

Figure 1.16: Process B- Extractive and Conventional Distillation Using NMP

1.6.7) Dimethylformamide (DMF) Solvent Extraction Process

This process, licensed by Nippon Zeon, consists of four sections: 1) first extractive distillation; 2) second extractive distillation; 3) butadiene purification; and 4) solvent purification. SeeFigure 1.7. In the first section of the plant, the hydrocarbon feed (C4 fraction) and the DMF solvent are fed to the first extractive distillation column, where the C4 stream is separated into two fractions:

1. the C4 raffinate 1 overhead product, which contains less soluble components (butane/butene); and, 2. The bottoms product, which contains the DMF solvent rich in butadiene/acetylene components which are the more soluble components.

43 At this stage, the C4 raffinate 1 product stream is available for downstream processing (MTBE, polyisobutylene, alkylation). The butadiene/acetylene rich solvent is fed to the first stripper column where the butadiene/acetylenes are stripped from the solvent and proceed overhead to the second extractive distillation section. The lean solvent from the stripper is cooled via heat recovery prior to being sent back to the extractor. In the second extractive distillation section, the butadiene/acetylene stream from the first section also is separated into two fractions, again using DMF as the solvent:[ Anonymous, (November 2006),Production of Butadiene]

1) butadiene/methyl acetylene rich overhead fraction; and, 2) bottoms fraction, containing the DMF solvent rich in vinyl acetylene which is more soluble in the DMF solvent than is butadiene or methyl acetylene.

The butadiene/methyl acetylene rich overhead fraction is sent on to the butadiene purification section of the process where the remaining acetylenes are removed using two distillation columns and a pure 1, 3-butadiene product stream is produced. The bottoms fraction from the second extractive distillation is fed to a stripper column where a vinyl acetylene rich stream is stripped from the solvent and used for fuel.

Since the DMF solvent is continuously circulated to the first and second extractive distillation columns, butadiene dimer, tar, and any water from the C4 feed stream tend to increase in concentration, thereby decreasing the effectiveness of the solvent. Therefore a part of the solvent is continuously passed to the solvent purification section of the plant to remove these impurities.

44

Figure 1.17: Process C- DMF Solvent Extraction Process

1.6.8) Aqueous Separation and Acetonitrile (ACN) Extraction

In this process, licensed by Shell, the hydrocarbon feed (C4 fraction) is routed to an extractive distillation system. The separation is achieved in an aqueous

solvent

environment,

where

the

top

product

contains

the

butanes/butylenes and the bottoms stream contains the butadiene and acetylenes. Acetonitrile (ACN) is used as the extraction solvent. As illustrated in Figure 1.8, the butadiene is then stripped from the extraction solvent and may be fed to a topping column where residual light ends (primarily methyl acetylene) are rejected. Heavier acetylenes such as vinyl acetylene and ethyl acetylene are rejected as a side stream from the solvent stripping operation. Bottoms product from the topping column can be fed to a postfractionator where residual olefins (cis-2/trans-2 butene) and remaining trace heavy ends (vinyl, ethyl and heavier acetylenes, 1,2-butadiene,  dimer,  C5’s  and  heavier)  are  rejected  to  the  bottoms.

45 The overhead butadiene product is chilled and passed through a coalescer to remove entrainedwater before being sent to the rundown tanks in the tank farm. Tertiary butyl catechol (TBC) is added to the butadiene to inhibit the formation of peroxides. It is also common to use an in process inhibitor that is removed prior to the addition of TBC. Hydrocarbon streams exiting the process can be washed with water for the removal of ACN.The recovered solvent can be concentrated and returned to the extraction section.[ Anonymous, (November 2006),Production of Butadiene]

Figure 1.18: Process D - Aqueous Separation and ACN Extraction Process

46 1.7)

JUSTIFICATION OF PROCESS CHOSEN

PRODUCTION OF CRUDE BUTADIENE 1 Temperature

2

1450-1525

°F 1100-1260 °F

3 930-1110

(790-830 °C)

(600-680 °C)

Pressure

98-118 kPa

15-20 kPa

Reaction

endothermic

endothermic

exothermic

% of butadiene

2-16%

30% - 50%

70-90%

Side products

hydrogen,

°F

(500-600 °C).

(exo/endo) _

ethylene, propylene,

fuel

gas, fuel oil Raw material

Ethane, propane, n-butane/n-Butenes butane,

n-Butenes

naphtha,

gas oil catalysts

chromium/alumina

bismuth molybdate

Table 1.1.5: Properties of Crude Butadiene

Among the choices for implementing the production of butadiene (polymer grade) from C4 fractions, we have chosen to do a coupling of catalytic dehydrogenation of n-butane and n-butene (the Houdry process) and oxidative dehydrogenation of n-butene (the Oxo-D or O-X-D process).

We have chosen this route because the coupling of the non oxidative catalytic dehydrogenation with the oxidative dehydrogenation of the n-butenes

47 formed provides a very much higher yield of butadiene based on n-butane used. The non oxidative dehydrogenation can also be operated in a gentler manner. This coupling process features particularly effective utilization of the raw materials. Thus, losses of the n-butane raw material are minimized by recycling unconverted n-butane into the dehydrogenation. The isomerization of 2-butene to 1-butene also yield 1-butene as the product of value. So because of the coupling process method, we get 1-butene as our byproduct after butadiene as a major production.

For our by product, 1-butene is a linear alpha olefin (alkene), produced either by separation from crude C4 refinery streams or from the reaction of ethylene. It is distilled to give a very high purity product.[ Anonymous, (2007),Butadiene Market Demand]

Butene-1 can be produced directly from C4 cracking and also by extraction from C4's mixtures out of ethylene crackers. It is used as a copolymer in polyethylene alkylates gasoline, polybutenes, butadiene; as intermediates for C4 and C5 aldehydes, alcohols and other derivatives; and in the production of maleic unhydride by catalytic oxidation.

Demand for 1-butene also has a big contribution in chemical area, like Japan’s   chemical   markets   for   1-butenes will grow strongly and for isobutylene very slowly during 2007–2012. Since August 2004, 1-butene demand for propylene production via butylene metathesis quadrupled to 2007, but will slow over the forecast period.[ Anonymous, (2007),Butadiene Market Demand]

While for application for 1-butene is used in the manufacture of a variety of other chemical products. It fills an important role in the production of materials such as linear low density polyethylene (LLDPE). The co-polymerisation of ethylene and 1-butene produces a form of polyethylene that is more flexible and more resilient. 1-butene can also help to create a more versatile range of

48 polypropylene resins. It is also used in the production of polybutene, butylene oxide and in the C4 solvents secondary butyl alcohol (SBA) and methyl ethyl ketone (MEK).

1.8)

SEPARATION

OF

BUTADIENE

BY

EXTRACTIVE

DISTILLATION

The C4 product gas stream is separated by means of extractive distillation into a stream that consisting substantially of n-butane and 2-butene and a product stream which is butadiene. To this end, C4 product gas stream is contacted in an extraction zone with an extractant, preferably an N-methyl-2pyrrolidone (NMP)/water mixture. Suitable extractants are butyrolactone, nitriles such as acetonitrile, propionitrile, methoxypropionitrile, ketones such as acetone, furfural, N-alkyl-substituted cyclic amides (lactams) such as N-alkylpyrrolidones, especially NMP. In general, alkyl-substituted lower aliphatic amides or N-alkylsubstituted cyclic arnides are used.[ Anonymous, (1996),Process of Butadiene production]

The effects of the solvents used will be taken into consideration in the comparison for the separation of C4 product to butadiene.

49 Solvent Hazard Identification methoxy-proprio-nitrile (MOPN)

Effect  Toxic fumes of acrylonitrile and methanol may be released.  Uncontrolled burning may also result in the release of highly toxic hydrogen cyanide (HCN) vapours.

n-methyl-2pyrrolidone (NMP)

 Stable, but decomposes upon exposure to light.  Combustible

Dimethylformamide (DMF)

 Hazardous  DMF has been linked to cancer in humans, and it is thought to cause birth defects

Acetonitrile (ACN)

 Toxic and flammable.  It is metabolized into hydrogen cyanide and thiocyanate

Table 1.1.6: Effects of the each solvents

It is noticed that all types of solvents used are hazardous and toxic. MOPN and ACN are unfavourable in terms of environmental friendly as both of these solvents will metabolize into hydrogen cyanide (HCN), which causes threat not only to human health, but also environment. Thus, NMP is the most acceptable solvent to be used in the process as it causes least harm to the environment.

Therefore Extractive distillation process by using aqueous n-methyl-2pyrrolidone (NMP) is the most environmental friendly method to be used as the solvent (NMP) causes less harm to the environment and provides the process with a safer mode due to the high autoignitable temperature.

50 1.9)

SITE SELECTION & PLAN LAYOUT

There are some factors that should be considered in selecting the suitable site. This is very important because the characteristics of a site location will have a market effect on the success or otherwise of a commercial venture. The choice of the final site should be based on a complete survey of the advantages of various geographical areas, in addition to the advantages and disadvantages of available industrial estates.

There are locations designated which have potential to fulfill all the criterion shows above. The locations is Kerteh Industrial Estate, Terengganu

1.9.1) Plant Location

The basic site selection process takes account the criteria of the economic geography of the area in which the plant site is placed. Several important criteria or factors should be considered, such as:

51

Figure 1.19: Map of Plant Location

1.9.1.1)

Land price

The first factor to consider is the land price. The cost of a land depends on the location of the property and may vary between rural district and a highly industrialized area. [49]

1.9.1.2)

Raw Materials

This is particularly important if large volumes of raw materials are required by plant. The availability of raw material is important due to its location from plant, transportation and its purity to ensure the least cost is needed.

52 1.9.1.3)

Transportation

Water, railroads and highways are the common means of transportation used by major industrial concerns. These types of connection needed to ensure the distribution of the product is optimize.

1.9.1.4)

Water Supply

The plant should be located where water supply is available for the purposes of cooling, washing, steam generation and as a raw material. Temperature, mineral content, purification (treatment) and cost of water ought to be considered.

1.9.1.5)

Energy supply

The location of a plant site should be near to hydroelectric installations if the plant is using electrolytic process. The local cost of power can help to determine whether power should be purchased of self-generated.

1.9.1.6)

Labors

Type and supply of labors available in the vicinity of a proposed plant site must be examined. Consideration should be given to prevailing pay scales and restrictions on number of working hours per week

53 1.9.1.7)

Waste management

The site selected for a plant should have adequate capacity and facilities for correct waste disposal. In choosing a plant site, the permissible tolerance levels for various methods of waste disposal should be considered carefully, and attention should be given to potential requirements for additional waste treatment facilities.[ Anonymous, (November 1996),Produce of Butadiene, Website]

1.9.1.8.1 Availability of raw materials

Location/ Characteristics

Kerteh Industrial Estate, Terengganu

Land Price

RM 0.18 -RM5.60 per ft 2

Sources of Raw Material (Natural gas

-Petronas LNG Sdn Bhd at Kerteh - Reserved natural gas until 2007 (53 trillion cubic feet)

Transportation Water Supply

East Coast highway Kemaman Port Terengganu Water Company (SATU

Electricity Supply

Paka power plant

Labor

- UMP, UMT, UTM, UiTM, TATIUC< POLISAS

Area Available

1184.2 hectares

Preferred Industry Type

Chemical

Nearest town

-15 km from Kemaman -160 km from Kuantan -220 km from Kuala Lumpur

Waste management

Monthly charges vary for disposal of toxic waste carried out by Kualiti Alam according to the type and quantity of waste.

Table 1.1.7: Characteristics of location

54 1.9.1.8.2 Selected Site  Industrial land is available in all districts in the form of developed, semi developed  as  well  as  raw  land.  Terengganu’s  industrial  land  is  among  the   cheapest in Malaysia, at RM0.18 - RM5.60 or (US$0.06 - $1.75 per square foot) compared to other states, where land sells from lows of RM2.00 RM4.50 psf to highs of RM18.00 - RM22.00 psf.  The location very near with Kemaman and Kuantan town provided local community factors such as banking, medical, security, housing and entertainment  Our plant needs the supplier to supply major raw material that is C4

 For delivering process of ammonia, Kerteh can be link to Kuala Lumpur by East Coast highway which can reduce traveling time from four hours to only two hours by truck also can be export by ships through Kemaman Port

 For the manpower or labor, UMP,TATIUC and POLISAS with quality educational level and other technical schools can provide skilled and semi skilled workers to operate and handle our plant

 Others universities such as UM, UTM, UTP, UPM, UKM and UiTM also can provide many graduate students to work in our plant

55  Terengganu employs an integrated environmentally friendly waste management system for industrial wastes that meets international best practices.[ Anonymous, (November 1996),Produce of Butadiene, Website]

55

N-Butane

N-Butane

Utilities Area

Butadiene

Butadiene

Waste treatment

Cooling Tower Butadiene Plant

Control Room

Loading / Unloading Area

Fire Water Tank Fire Station Mosque Cafeteria

Fire Water Tank

Cafeteria Admininstration Building

Master Evacuation Area

Workshop

Laboratory

Mosque

Post Guard

Figure 1.20: Plant Layout

56

CHAPTER 2

PROCESS SYNTHESIS AND FLOWSHEETING

2.1 SYNTHESIS OF PROCESS FLOW DIAGRAM 2.1.1 Introduction Process synthesis and flow sheeting is the most important in plant design. In this chapter, it will cover material and energy balance and simulation by using ASPEN. Generally, the processes that choose in producing butadiene from butane is catalytic dehydrogenation and oxidative dehydrogenation. This route absolutely gives an optimum purity of butadiene. Here, 50 000 MT/yr butadiene from butane will produced.

2.1.2

Process Flow Diagram (PFD) Process flow diagram in figure 2.1 below described details flow of

butadiene process. Basically for the PFD diagram, the major and minor include together except the controller system that explained details in process safety studies. Catalytic dehydrogenation process used 1 basic chemical which is butane and water and air used as a combustion process to remove hydrogen. Then produce 1-butene and 2-butene. For second reaction, used Oxidative

57 dehydrogenation process to produce 1-butane, 2-butene to butadiene from hydrogen removed. To produce butadiene, there are 6 equipments involved. The equipments are:

a. The Fixed Bed Reactor – Reaction occurs here b. The NMP solvent recovery system that function to recover the NMP solvent. c. The distillation for separate light and heavy end which is 1-butane and butane d. Heat transfer for maintain the temperature in the process e. Absorption for separate inert gas and C4 hydrocarbon f. Extractive distillation for separate butadiene and C4 hydrocarbon

58

C-102A/B Air Pump Compressor

C-103A/B Steam Compressor

R-101 Fixed Bed Ractor (BDH)

C-104A/B Compressor Stream 5

E-101 Heat Exchanger

C-105A/B Compressor Stream 6

R-102 Fixed Bed Reactor (Oxo-D)

TK-2 Storage Tank (1-Butene)

P-107A/B Pump Stream 14

P-106A/B Pump

E-108 D3 Reboiler

P-105A/B Pump

E-107 D3 Heat Exchanger

P-104A/B Pump Stream 12b

R-103 Fixed Bed Reactor (Isomerization)

C-106A/B Compressor Stream 7

P-101A/B Pump Stream 8

C-107A/B Compressor Stream 9

T-101 Absorption Tower

P-102A/B Pump Stream 11

D-102 Solvent Recovery Column

E-104 D1 Reboiler

E-103 D1 Heat Exchanger

E-105 D2 Reboiler

E-106 D2 Heat Exchanger

D-101 Extractive Distillation Column

TK-1 Storage Tank (Butadiene)

P-103A/B Pump Stream 12

E-102 Condenser (Air Cooled)

8b

C-101A/B Butane Feed Compressor

Calcium Oxide

6

Water

8

Air

P-101A/B

8a

C-105A/B

Air

3

3b

C-103A/B

C-102A/B

n-Butane

7

5

2

C-106A/B

C-104A/B

1

7a

9

E-102

E-101

9a

C-107A/B

P-102A/B

4

T-101 R-102

R-101

C-101A/B

CW

CW 10

CW

11

E-106

E-107

E-103

14

13 15

12a

TK-2 1-Butene

P-106A/B Reflux Drum

TK-1 Butadiene

Reflux Drum

12b

NMP 2,000 kg/h

Reflux Drum 16

16a

P-105A/B P-104A/B

E-108

E-104

E-105

Steam R-103

D-102

D-101

Steam

P-107A/B

12

P-103A/B NMP Recovery

Figure 2.1.: Process Flow Diagram

16b

D-103

Steam

12a

Steam

95

2.2

Manual Calculation of Materials and Energy Balances

2.2.1 Mass Balance Plant complete if the amount that targeted is achieved. Here, in order to achieve the target, the amount needed for each hour need to be known operation day are assumed 333 days with 32 days shutdown and the plant operation 24 hours per day.

Butadiene will produce 50,000 MT/ year: 50,000 MT

1000 kg 1 year

year

1 MT

1 day

= 6256.2563 kg/hr

333 days 24 hours

Therefore, to make sure we could achieve the amount production, the mass flow rate production must be 6256.2563kg/hr. So below is the assumption that our group have been made.

Assumptions: 

The processes follow the law of conservation of mass where:



Material out = Material in + Generation – Consumption – Accumulation



Steady-state condition in all equipment



Pure reactants are used



No leakage in pipes and vessel in the system.



All stream flow rate is in unit kg/h.



Catalyst used in reactor does not contribute in mass.



The entire components in the system behave as ideal condition.



The total input of any substance to a pump, valve or mixer is assumed equal to the total output of the substance where no reaction occurs in that device.

60

Components involve: Components

Formula C4 H10

Molecular Weight 58.12

1-Butene 2-butene 1,3-Butadiene

C4H8 C4H8 C4H6

56.10 56.10 54.09

Water Carbon dioxide Hydrogen

H2O CO2 H2

18.016 44.01 2.016

Oxygen

O2

32.00

Nitrogen

N2

28.02

Butane

Table 2.1.0: Components formulas and Molecular Weight

2.2.2 Energy Balance 2.2.2.1 Enthalpies for Vapor Mixtures

Methods suggested by Biegler et al. in calculating enthalpies for vapor mixtures are as below:

H v (T , y)  H f  H T   k y k H of .k (T1 )   y k  C po.k (T )dT T2

k

where;

T1

Hf.k(T1) is the heat of formation for component k at

T1 Cp is heat capacities for component k.

61 2.2.2.2 Enthalpies for Liquid Mixtures Enthalpies for liquid mixtures are evaluated from the ideal vapor enthalpy and subtracting the heat of vaporization at the saturated condition.

k H L (T , x)  H f  H T   k xk H of .k (T1 )   xk  C po.k (T )dT  H vap (T ) T2

T1

k

where,

Hf.k(T1) is the heat of formation for component k at T1 Cp is heat capacities for component k. Hvap is heat of vaporization

For a mixed stream, both equations are applied based on the vapor/liquid fraction involved. If there is no reaction occur in a unit (i.e. initial component = final component), enthalpy change for the unit is express as below:

Qv  n y k [  C po.k (T )dT  H f ,k ] (vapor) T2

k

T1

k QL  n xk [  C po.k (T )dT  H f ,k (T2 )  H vap (T1 )] (liquid) T2

k

T1

2.2.2.3 Heat of Vaporization Hvap is heat of vaporization of specific component at specific temperature. It could be found through the Watson Method: T k T

k k H vap (T )  H vap (Tb )[ T ck T ]0.38 c

b

where, Tck is critical temperature for component k.

62 Tb is the boiling temperature for component k

Component

Heat Of Vaporization (kJ/mole) Butane 22.305 1-Butene 21.916 2-butene 21.916 1,3-Butadiene Water 40.706 Carbon dioxide 15.326 Hydrogen 0.964 Oxygen 6.82 Nitrogen 5.577 (Source: Chemical Engineering, Volume 6) Table 2.1.1: The Heat Of Vaporization for Each Component at boiling point

2.2.2.4 Heat of Formation

Component Heat Of Formation (kJ/mole) Butane -126.9 1-Butene -0.12558 2-butene -11.17662 1,3-Butadiene 110,21738 Water -241.9508 Carbon dioxide -393.60958 Hydrogen 0 Oxygen 0 Nitrogen 0 (Source: Chemical Engineering, Volume 6) Table 2.1.2: The Heat of Formation for Each Component

63 2.2.2.5 Heat Capacity CP = A + BT + CT2 + DT3

Components Butane 1-Butene 2-butene 1,3-Butadiene Water Carbon dioxide Hydrogen Oxygen

A 92.30x10-3

B 27.88x10-5

C -15.47x10-8

D 34.98x10-12

33.46x10-3 36.11 x10-3

0.6880x10-5 4.233 x10-5

0.7604x10-8 -2.887 x10-8

28.84 x10-3

0.00765 x10-5

29.10 x10-3

1.158 x10-5

0.3288 x10-8 -0.6076 x10-

-3.593x10-12 7.464 x10-12 -0.8698 x10-

8

Nitrogen

12

1.311 x10-12 -2.871 x10-

12 29.00 x10-3 0.2199 x10-5 0.5723 x10-8 (Source: Perry Handbook) Table2.1.3: The Parameters of Heat Capacity, Cp (J/mole K) or (J/moleC)

2.3

Mass and Energy Balance

2.3.1. Mass Balance in Fix Bed Reactor (FBR)-BDH reaction R-101 Stream 3b 2 4 5

Description Inlet Inlet Inlet Outlet

Temperature(0C) 200 300 30 550

Pressure(atm) 1.0 2.0 4.5 1.0

Table2.1.4: The Operation Condition of Reactor 1

64

Stream 3b

Stream 2

Stream 5

Stream 1

Function

To be a reaction place among butane, air and steam and create combustion, then remove hydrogen in dehydrogenation process.

Condition Phase

: Gases phase

Operating Temperature

: 550 oC

Operating Pressure

: 1.0 atm

Catalyst

: Chromium alumina

Promoter

: 1-butene and 2-butene

Conversion

: 80 %

65

Stream Name Stream Phase Pressure Temperature oC

S4

S2

S3b

S5

Gas

Gas

Gas

Gas

1.0 atm

2.0 atm

1.0 atm

1.0 atm

30oC

300oC

200oC

550oC

6885.37

15084.34

5341.72

-

-

-

Total Flow rate(Kg/hr) For Initial Feed Total Flow rate (Kg/hr) Outlet or

37018.14

recycle back Table 2.1.5: Input stream of FBR (1st reactor)

Input Stream (kg/kg) Component

Output

S4

S2

S3b

S5

1.0

0

0

0.1860

1-Butene

0

0

0

0.0614

2-butene

0

0

0

0.1316

1,3-Butadiene

0

0

0

0.0127

Water

0

0

1.0

0.1443

Carbon dioxide

0

0

0

0.0065

Hydrogen

0

0.1429

0

0.2245

Oxygen

0

0.2853

0

0

Nitrogen

0

0.5718

0

0.2330

Butane

Table 2.1.6: Composition in FBR

66

Input Stream (kg/hr)

Output

Component

S1

S2

S3b

S5

Butane 1-Butene 2-butene 1,3-Butadiene Water Carbon dioxide Hydrogen Oxygen Nitrogen

6885.373 0 0 0 0 0 0 0 0

0 0 0 0 0 0

0 0 0 0 5341.72 0 0 0 0

6885.37 2269.21 4871.58 470.13 5341.72 240.62 8310.57 0 8625.23

TOTAL

6885.37

5341.72

37018.14

15084.34 15084.11

Table 2.1.7: Mass Balance in FBR (1st reactor)

2.3.2 Energy Balance in FBR (1st reactor)

Catalytic Dehydrogenation of n-butane & n-butene reaction (the Houdry process) Butane  1-butene + Hydrogen Butane  2-butene + Hydrogen C4H10  C4H8-1 + H2 C4H10  C4H8-2 + H2

67 Input Stream ΔH1 (25 oC) Inlet FBR

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole flow rate 0 0 0 0 0 0 1077.78 134.49 308.04

Stream 2 ΔH,  Enthalpy   (kj/mol) 0 0 0 0 0 0 7894.647 8459.949 8206.952

Q=  nΔH  (J) 0 0 0 0 0 0 8508693 1137779 2528069 12174541

Table 2.1.8: Heat of FBR input

Inlet FBR

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 0 0 0 0 296.76 0 0 0 0

Stream 3b ΔH,  Enthalpy   (kj/mol) 0 0 0 0 6009.047 0 0 0 0

Table 2.1.9: Heat of FBR input

Q=  nΔH  (J) 0 0 0 0 1783258 0 0 0 0 1783258

68 Inlet FBR

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 118.71 0 0 0 0 0 0 0 0

Stream 1 ΔH,  Enthalpy   (kj/mol) 490.207 0 0 0 0 0 0 0 0

Q=  nΔH  (J) 58194.08 0 0 0 0 0 0 0 0 58194.08

Table 2.1.10: Heat of FBR input

Output Stream ΔH2 (25oC

Component C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2

550oC)

n, Mole Rate 118.7133 40.52161 86.9925 8.706111 296.7622 5.468636 259.7053 0 4312.615

Outlet FBR S5 ΔH,  Enthalpy  (kj/mol) 62884.548 54976.598 54709.397 50283.658 15107.457 18599.423 12328.319 0 12849.743

Q=  nΔH 7465230.69 2227740.11 4759307.22 437775.114 4483322.51 101713.481 3201729.94 0 55415994.4

TOTAL 78092813.5 Table 2.1.11: Heat of FBR output

69 ΔHf ((25oC

550oC)

Component C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Heat formation (endothermic reaction) n, Mole Rate ΔH,  Enthalpy  (kj/mol) Q=  nΔH 118.7133 -1.2579E+05 -1.49E+07 40.52161 -5.4E+02 -2.19E+04 86.9925 -5.4E+02 -4.70E+04 8.706111 1.0924E+05 9.51E+05 296.7622 -2.42E+04 -7.18E+06 5.468636 -3.93509E+05 -2.15E+06 259.7053 0 0 0 0 0 4312.615 0 0 -2.34E+07 Table 2.1.12: Heat formation of process

Q For FBR (1st reactor) =  nΔH2 + nΔHf - nΔH1 = [(H5+Hf)-(H2+H3b+H1] = [(78092813.5+ -2.34E+07)-(12174541+1783258+58194.08)] = 6.4E+7 kJ/hr

2.3.3 Mass Balance in Fixed Bed Reactor (2nd reactor) - Oxo-D reaction R-102 Stream 5 6 7

Description Inlet Inlet Outlet

Temperature(0C) 550 34 450

Pressure(atm) 1.0 1.0 2.0

Table2.1.13: The Operation Condition of Reactor 2

70 Stream 6

Stream 5

Stream 7

Function To be a reaction place among 1-butene,2-butene and Oxygen to crack or remove the hydrogen in oxidative dehydrogenation process

Condition Phase

: Gases phase

Operating Temperature

: 450 oC

Operating Pressure

:

Catalyst

: Bismotmolybidat

Promoter

: Butadiene

1 atm

71 Component

Mass Flow rate (kg/hr) Inlet Stream

Outlet Stream

S5

S6

S7

C4H10

6885.374

0

6885.373

C4H8-1

2272.914

0

0

C4H8-2

4871.587

0

3133.268

C4H6

470.1304

0

4292.773

H2O

5341.718

0

9797.163

CO2

240.6179

0

885.9137

H2

8310.572

0

2605.628

O2

0

2605.63

29248.18

N2 TOTAL

8625.227

9802.12

8292.412

37018.14

12407.75

65140.71

Table 2.1.14: Inlet and Outlet Stream of FBR (2nd reactor)

2.3.4

Energy Balance in FBR

COMPONENT Mole Rate 118.7133 40.58775 86.99263 8.706119 296.7621 5.468589 259.7054 0 4312.614

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2

Inlet S5 ΔH,  Enthalpy  (kj/mol) 62884.548 54976.598 54709.397 50283.658 15107.457 18599.423 12328.319 0 12849.743

TOTAL Table 2.1.15: Inlet Stream of FBR (2nd reactor)

Q=  nΔH 7465235.03 2231376.42 4759314.06 437775.486 4483320.83 101712.593 3201730.71 0 55415975.1 78096440.3

72

COMPONENT Mole Rate 0 0 0 0 0 0 0 93.05821 4901.06

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2

Inlet S6 ΔH,  Enthalpy  (kj/mol) 0 0 0 0 0 0 0 270.038 262.032

Q=  nΔH 0 0 0 0 0 0 0 25129.2541 1284234.55 165968.1956

TOTAL Table 2.1.16: Inlet Stream of FBR (2nd reactor)

COMPONENT Mole Rate 118.7133 0 55.95121 79.4958 544.2868 20.1344 81.42588 1044.578 4146.206

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2

Outlet S7 ΔH,  Enthalpy  (kJ/mol) 62884.548 0 54709.397 50283.658 15107.457 18599.423 12328.319 13249.802 12849.743

Q=  nΔH 7465233.947 0 3061057.195 3997339.433 8222789.93 374488.2647 1003844.162 13840449.78 53277681.53 91242884.24

TOTAL Table 2.1.17: Outlet Stream of FBR (2nd reactor)

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

n,Mole Rate 118.7133 0 55.95121 79.4958 544.2868 20.1344 81.42588 1044.578 4146.206

Hf -1.2579E+05 -5.4E+02 -5.4E+02 1.0924E+05 -2.42E+04 -3.93509E+05 0 0 0

Table 2.18: Heat of Formation for the process

Q=nΔH -1.49E+07 0 -3.02E+04 8.68E+06 -1.32E+07 -7.92E+06 0 0 0 -2.7E+07

73

Q For Condenser (air cooled) = nΔH2 + nΔHf - nΔH1 = [(H7+Hf)-(H5+H6)] = [(91242884.24-27000) – (78096440.3+165968.1956)] = 1.3E+7 kJ/hr

2.3.5 Mass Balance in Heat Exchanger E-101 Stream 7a 8 9a

Description Inlet Outlet Outlet

Temperature(0C) 200 40 40

Pressure(atm) 2.0 1.0 0.95

Table2.1.19: The Operation Condition of Heat Exchanger

Stream, 7a Stream 7

E-101

Function To remove the water, maintain the temperature and recycle the water into condenser

Condition Phase

: Liquid phase

Operating Temperature

: 160 oC

Operating Pressure

: 1 atm

74

Component C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mass Flow rate(kg/hr) Inlet Outlet S7a S8 S9a 10052.60487 1.45536686 10052.81368 0 0 0 4574.5534 994.0155654 3576.481791 6267.423474 2.91073372 6274.953413 1430.380107 13534.9118 733.0176645 1293.42882 16.00903546 1272.711989 12106.87418 0 12106.87417 3804.202412 1.45536686 3818.135967 42702.17207 2.91073372 42716.40302

Table 2.1.20: Inlet and Outlet Streams

2.3.6 Energy balance for Heat Exchanger

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.9539919 0 81.53266795 115.8681384 79.39939533 29.38943013 6005.393937 118.8850405 1524.369831

Inlet Stream 7 ΔH,  Enthalpy  (kj/mol) 21056.087 0 18566.496 16781.533 6009.047 7258.533 4988.39 5334.38 5175.577

Q=  nΔH 3641734.301 0 1513775.953 1944444.988 477114.6983 213324.1485 29957247.06 634177.9825 7889493.437 46271312.57

Table 2.1.21: Energy balance at inlet stream for Heat Exchanger

75

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 0.025039431 0 17.71642692 0.053811793 751.3134498 0.363759042 0 0.045481636 0.103906533

Outlet Stream 7a ΔH,  Enthalpy  (kj/mol) 1491.535 0 1340.899 1220.147 1125.511 598.945 0 450.326 436.974

Q= nΔH 37.34718802 0 23755.93914 65.65829835 845611.5522 217.8716597 0 20.48156307 45.40445352 869754.2545

Table 2.1.22: Energy balance at Outlet stream for Heat Exchanger

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.9575845 0 63.7439498 116.0073471 40.68929584 28.91870004 6005.393935 119.3204777 1524.877843

Outlet S9a ΔH,  Enthalpy  (kj/mol) 1491.535 0 1340.899 1220.147 1125.511 598.945 422.643 450.326 436.974

Q=  nΔH 257972.2908 0 85474.19854 141546.0166 45796.25004 17320.7108 2538137.709 53733.11345 666331.9706 3806312.26

Table 2.1.23: Energy Balance at Outlet stream for Heat Exchanger

76

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.9575845 0 63.7439498 116.0073471 40.68929584 28.91870004 6005.393935 119.3204777 1524.877843

Outlet Heat Vaporization ΔH,  Enthalpy  (kj/mol) 22.305 21.916 21.916 0 40.706 15.326 0.964 6.82 5.577

Q=  nΔH 3857.8189 0 1397.012404 0 1656.2985 443.2079968 5789.199753 813.7656579 8504.24373 22461.54694

Table 2.1.24: Heat Vaporization at Heat Exchanger Q For Heat Exchanger = nΔH2 + nΔHv - nΔH1 = [(H8+H9a+Hv)-(H7a)] =[ (869754.2545+ 3806312.26+ 22461.54694)46271312.57] = -41572784.51 kJ/hr

2.3.7 Mass Balance in Condenser (Air cooled) E-102 Stream 7a 8 9

Description Inlet Outlet Outlet

Temperature(0C) 200 30 40

Pressure(atm) 1.0 1.0 1.0

Table2.1.25: The Operation Condition of Condenser (air cooled)

77

Stream 8

Stream 7a

Stream 9

E-102

Function To recycle back the water to cover another process

Condition Phase

: Liquid phase

Operating Temperature

: 100 oC

Operating Pressure

: 1 atm

Component

Mass Flow rate (kg/hr) Inlet Stream

Outlet Stream

S7a

S8

S9

C4H10

8827.6175

1.2811

8823.9913

C4H8-1

0

0

0

C4H8-2

4017.1089

874.9719

3139.3046

C4H6

5503.6896

2.5621

5507.9241

H2O

12560.7726

11913.9664

643.4160

CO2

1135.8145

14.0918

1117.1399

H2

10631.5582

1.2811

3351.4198

O2

3340.6310

1.2811

37494.8927

N2 TOTAL

37498.5832

1.2811

10626.9703

∑83515.7755

∑83515.7755

Table 2.1.26: Inlet and Outlet Stream of Condenser (air cooled)

78 2.3.8 Energy Balance in Condenser (air cooled)

COMPONENT Mole Rate 152.2003 0 71.7341 101.9202 697.8207 25.8140 5315.7791 106900.192 1339.2351

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2

Inlet S7a ΔH,  Enthalpy  (kJ/mol) 497798.8 0 48081.6 46364.4 15454.8 37778.4 1717.2 27475.2 24040.8

TOTAL

Q=  nΔH 75765126.7 0 3449090.303 4725468.921 10784679.35 975211.6176 9128255.871 2937104155 32196283.19 137024116

Table 2.1.27: Inlet Stream of condenser (air cooled)

COMPONENT Mole Rate 0.0221 0 15.6245 0.04745 661.887 0.3203 5315.7791 104.3947 1339.2351

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2

Outlet S8 ΔH,  Enthalpy  (kJ/mol) 7540 0 874.972 7020 11913.966 5720 260 4160 3640

TOTAL Table 2.1.28: Inlet Stream of Condenser

Q=  nΔH 166.634 0 13671 333.099 7885699.214 1832.116 1382102.566 1806612920 4874815.764 1820771540

79

COMPONENT Mole Rate 152.1378 0 56.0590 101.9986 35.7453 25.3895 1675.7099 1171.7154 379.5347

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2

Outlet S9 ΔH,  Enthalpy  (kj/mol) 10086.2 0 9738.4 9390.6 3130.2 7651.6 347.8 5564.8 4869.2

Q=  nΔH 1534492.278 0 545924.9656 957828.0532 111889.9381 194270.2982 582811.9032 6520361.858 1848030.361 12295609.66

TOTAL Table 2.1.29: Outlet Stream Condenser (air cooled)

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

n,Mole Rate 152.1599 0 56.1065 102.0461 697.6323 25.7098 6991.489 1276.1101 1718.7698

Hv 22.305 0 21.916 171.005 40.656 304.2 0 0 0

Q=nΔH 3393.9266 0 1229.6301 17450.3933 28362.9388 7820.9212 0 0 0 58257.81

Table 2.1.30: Heat of Formation for the process

Q For Condenser(cooled) = nΔH2 + nΔHv - nΔH1 = [(H9+H8+Hv)-(H7a)] = [(12295609.66+1820771540+58257.81)–(137024116)] = 1696101291 kJ/h

80 2.3.9 Mass Balance in Compressor C-107A/B Stream 9 9a

Description Outlet Inlet

Temperature(0C) 100 40

Pressure(atm) 1.50 0.95

Table2.1.31: The Operation Condition of Compressor

Stream 9a

Stream 9

E-107

Mass Flow rate(kg/hr) Component C4H10 C4H8-1 C4H8-2 C4H6 H2 O CO2 H2 O2 N2 TOTAL

Inlet S9 8823.9913 0 3139.3046 5507.9241 643.4160 1117.1399 3351.4198 37494.8927 10626.97 ∑70705.0584

Outlet S9a 8823.9913 0 3139.3046 5507.9241 643.4160 1117.1399 3351.4198 37494.8927 10626.97 ∑70705.0584

Table 2.1.32: Mass Balance Inlet and Outlet Stream at Compressor

81 2.4.0 Energy balance for Compressor

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.9575845 0 63.7439498 116.0073471 40.68929584 28.91870004 6005.393935 119.3204777 1524.877843

Inlet S9 ΔH,  Enthalpy  (kJ/mol) 1491.535 0 1340.899 1220.147 1125.511 598.945 422.643 450.326 436.974

Q=  nΔH 257972.2908 0 85474.19854 141546.0166 45796.25004 17320.7108 2538137.709 53733.11345 666331.9706 3806312.26

Table 2.1.33: Energy Balance at Compressor

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.9575845 0 63.7439498 116.0073471 40.68929584 28.91870004 6005.393935 119.3204777 1524.877843

Outlet S9a ΔH,  Enthalpy  (j/mol) 2122.514 0 7189.784 6609.026 2543.847 3038.834 2122.514 2264.693 2197.475

Table 2.1.34: Energy Balance at Compressor

Q=  nΔH 367104.8945 0 458305.2304 766695.5734 103507.3431 87879.12892 12746532.7 270224.2507 3350880.938 18151130.06

82 COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2

n, MOLE RATE 172.9575845 0 63.7439498 116.0073471 40.68929584 28.91870004 6005.393935 119.3204777 1524.877843

Heat Vaporization, Hv 22.305 21.916 21.916 0 40.706 15.326 0.964 6.82 5.577

TOTAL

Q=  nΔH 3857.818922 0 1397.012404 0 1656.298476 443.2079968 5789.199753 813.7656579 8504.24373 22461.54694

Table 2.1.35: Heat Vaporization at Compressor

Q For Compressor = nΔH2 + nΔHv - nΔH1 =

H9+Hv-H9a

= [(18151130.06+ 22461.54694)-3806312.26] = 14367279.35 kJ/hr 2.4.1 Mass Balance in Absorption Tower T-101 Stream 9 10 11

Description Inlet Outlet Outlet

Temperature(0C) 100 40 40

Pressure(atm) 1.50 0.95 0.95

Table2.1.36: The Operation Condition of Absorption Tower

83

Stream 10

Stream 9

Stream 11

Function To remove the water and inert gases Condition Phase

: Gas phase

Operating Temperature

: 60 oC

Operating Pressure

: 1 atm

Mass Flow rate(kg/hr) Component C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Inlet S9 8823.9913 0 3139.3046 5507.9241 643.4160 1117.1399 10626.9703 3351.4198 37494.8927 ∑70705.059

Outlet S10 S11 31.4889 8792.8990 0 0 41.9851 3099.8386 20.9926 5485.3111 152.1962 493.8602 787.2215 326.2029 10627.4904 0 3332.5711 16.4013 37487.4883 9.1118 ∑70705.059

Table 2.1.37: Inlet and Outlet Streams

84 2.4.2 Energy balance for Absorption Tower

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.9575845 0 63.7439498 116.0073471 40.68929584 28.91870004 6005.393935 119.3204777 1524.877843

Input Stream 9 ΔH,  Enthalpy  (kJ/mol) 2122.514 0 7189.784 6609.026 2543.847 3038.834 2122.514 2264.693 2197.475

Q=  nΔH 367104.8945 0 458305.2304 766695.5734 103507.3431 87879.12892 12746532.7 270224.2507 3350880.938 18151130.06

Table 2.1.38: Energy balance for Input Absorption Tower

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 0.617180174 0 0.852475051 0.442123622 9.624373527 20.37739335 6005.417956 118.6440757 1524.081355

Output Stream 10 ΔH,  Enthalpy  (j/mol) 1491.535 0 1340.899 1220.147 1125.511 598.945 422.643 450.326 436.974

Q=  nΔH 920.545831 0 1143.082943 539.4558111 10832.33827 12204.93786 2538147.861 53428.51203 665983.9262 3283200.66

Table 2.1.39: Energy Balance for Output Absorption Tower

85

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.3704507 0 62.95073242 115.5460243 31.23550603 8.445301996 0 0.584009114 0.296492623

Output Stream 11 ΔH,  Enthalpy  (j/mol) 1491.535 0 1340.899 1220.147 1125.511 598.945 0 450.326 436.974

Q=  nΔH 257096.5602 0 84410.57415 140983.1349 35155.90563 5058.271404 0 262.9944883 129.5595673 523097.0003

Table 2.1.40: Energy Balance for Output Absorption Tower COMPONENT

n, MOLE RATE

Heat Vaporization, Hv

Q=  nΔH

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

10018.68771 0 3531.976744 6250 562.7076412 371.6777408 0 18.68770764 8.30564784

22.305 21.916 21.916 0 40.706 15.326 0.964 6.82 5.577

223466.8294 0 77406.80232 0 22905.57724 5696.333056 0 127.4501661 46.320598 329649.3128

Table 2.1.41: Heat Vaporization at Absorption Tower Q For Absorption Tower = nΔH2 + nΔHv - nΔH1 = (H10+H11+HV) -H9 = (3283200.66+523097.0003+329649.3128)- 18151130.06 = -14015183.09 kJ/hr

86 2.4.3 Mass Balance in Extractive Distillation Column D-101 Stream 11 12 13

Description Inlet Outlet Outlet

Temperature(0C) 40 70 50

Pressure(atm) 1.5 1.0 1.0

Table2.1.42: The Operation Condition of Extractive Distillation Column

Stream 13 Reflux ratio

Stream 11

D-101

Stream 12

87 Function To remove the water, maintain the temperature and recycle the water into waste boiler Condition Phase

: Liquid phase

Operating Temperature

: 30

Operating Pressure

: 5 atm

o

C

Mass Flow rate(kg/hr) Component C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Inlet S11 8792.8990 0 3099.8386 5485.3111 493.8602 326.2029 0 16.4013 9.1118 ∑18223.6249

Outlet S12 S13 8035.3175 63.1945 0 0 3072.4673 0 23.8083 6256.2563 452.3586 0 297.6044 0 0 0 15.4754 0 7.1425 0 ∑18223.6249

Table 2.1.43: Inlet and Outlet Streams

88 2.4.4 Energy balance for Extractive Distillation Column

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.3704507 0 62.95073242 115.5460243 31.23550603 8.445301996 0 0.584009114 0.296492623

Input Stream 11 ΔH,  Enthalpy  (j/mol) 1491.535 0 1340.899 1220.147 1125.511 598.945 0 450.326 436.974

Q=  nΔH 257096.5602 0 84410.57415 140983.1349 35155.90563 5058.271404 0 262.9944883 129.5595673 523097.0003

Table 2.1.44 Energy balance for Input Extractive Distillation Column

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.352517 0 62.99014602 0 31.34050805 8.44265989 0 0.589654535 0.310872515

Output Stream 12 ΔH,  Enthalpy  (j/mol) 4659.945 0 4169.698 0 1521.028 1810.12 0 1354.907 1314.716

Q=  nΔH 803153.25 0 262649.8859 0 47669.79027 15282.22752 0 798.9270577 408.7090693 1129962.79

Table 2.1.45: Energy Balance for Output Extractive Distillation Column

89

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 1.0896 0 0 115.8566 0 0 0 0 0

Output Stream 13 ΔH,  Enthalpy  (kJ/mol) 209.3 0 0 7251.714 0 0 0 0 0

Q=  nΔH 228.0533 0 0 84078.7553 0 0 0 0 0 84306.8086

Table 2.1.46: Energy Balance for Output Extractive Distillation Column COMPONENT

n, MOLE RATE

Heat Vaporization, Hv

Q=  nΔH

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

172.352517 0 62.99014602 0 31.34050805 8.44265989 0 0.589654535 0.310872515

22.305 21.916 21.916 0 40.706 15.326 0.964 6.82 5.577

3844.322892 0 1380.49204 0 1275.746721 129.3922055 0 4.021443929 1.733736016 6635.709038

Table 2.1.47: Heat Vaporization at Extractive Distillation Column Q For Extractive Distillation Column = nΔH2 + nΔHv - nΔH1 = (H13+H12+HV)-H11 =( 83853.36747+1129962.79+6635.709038)- 523097.0003 = 697354.8662 kJ/hr

90 2.4.5 Mass Balance in Solvent Recovery Column D-102 Stream 12 12a 12c

Temperature(0C) 70 400 400

Description Inlet Outlet Outlet

Pressure(atm) 1.0 3.0 3.0

Table2.1.48: The Operation Condition of Recovery Column

Stream 12a

Stream 12

Function To recovery the NMP solvent Condition Phase

: Liquid phase

Operating Temperature

: 400

Operating Pressure

: 2 atm

o

C

91

Mass Flow rate(kg/hr) Component C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 NMP TOTAL

Inlet S12 10017.64535 0 3534.188123 0 564.5992524 371.5614618 0 18.86835548 8.70847176 0 ∑14515.57

Outlet S12a 7368.818521 6158.340946 452.8405315 0 37.73671096 480.4173588 0 13.06270764 5.80564784 0 ∑14515.57

S12c 0 0 0 0 0 0 0 0 0 2000 kg/hr 2000

Table 2.1.50: Inlet and Outlet Streams

2.4.6 Energy balance for Solvent Recovery Column

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.352517 0 62.99014602 0 31.34050805 8.44265989 0 0.589654535 0.310872515

Input Stream 12 ΔH,  Enthalpy  (j/mol) 4659.945 0 4169.698 0 1521.028 1810.12 0 1354.907 1314.716

Q=  nΔH 803153.25 0 262649.8859 0 47669.79027 15282.22752 0 798.9270577 408.7090693 1129962.79

Table 2.1.51: Energy balance for Input Solvent Recovery Column

92

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 126.7797347 109.7606528 8.071016656 0 2.094738327 10.9160954 0 0.408222371 0.207248343

Output Stream 12a ΔH,  Enthalpy  (j/mol) 53583.913 46892.33 46698.873 0 13233.507 16244.64 0 11639.901 11289.699

Q=  nΔH 6793354.275 5146932.752 376907.3818 0 27720.73431 177328.0401 0 4751.667981 2339.771414 12529334.62

Table 2.1.52: Energy Balance for Output Solvent Recovery Column COMPONENT

n, MOLE RATE

Heat Vaporization, Hv

Q=  nΔH

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

126.7797347 109.7606528 8.071016656 0 2.094738327 10.9160954 0 0.408222371 0.207248343

22.305 21.916 21.916 0 40.706 15.326 0.964 6.82 5.577

2827.821982 2405.514467 176.884401 0 85.26841834 167.3000781 0 2.784085702 1.155824494 5666.729257

Table 2.1.53: Heat Vaporization at Solvent Recovery Column

Q For Solvent Recovery Column = nΔH2 + nΔHv - nΔH1 = (H12a+HV)-H12 =[(12529334.62+5666.729257)-1129962.79] = 11405038.56 kJ/hr

93 2.4.7 Mass Balance in Isomerization Stage in FBR R-103 Stream 12 12a

Temperature(0C) 70 400

Description Inlet Outlet

Pressure(atm) 1.0 3.0

Table2.1.54: The Operation Condition of Isomerazation Stage in FBR

Stream 12a

Stream 12

Function To isomerizes 2-butene to 1-butene Condition Phase

: Liquid phase

Operating Temperature

: 40

Operating Pressure

: 2 atm

o

C

94

Mass Flow rate(kg/hr) Component C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Inlet S12 10017.64535 0 3534.188123 0 564.5992524 371.5614618 0 18.86835548 8.70847176 ∑14515.57

Outlet S12a 7368.818521 6158.340946 452.8405315 0 37.73671096 480.4173588 0 13.06270764 5.80564784 ∑14515.57

Table 2.1.55: Inlet and Outlet Streams 2.4.8 Energy balance for Isomerization Stage in FBR

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 172.352517 0 62.99014602 0 31.34050805 8.44265989 0 0.589654535 0.310872515

Input Stream 12 ΔH,  Enthalpy  (j/mol) 4659.945 0 4169.698 0 1521.028 1810.12 0 1354.907 1314.716

Q=  nΔH 803153.25 0 262649.8859 0 47669.79027 15282.22752 0 798.9270577 408.7090693 1129962.79

Table 2.1.56: Energy balance for Input Isomerization Stage in FBR

95

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 126.7797347 109.7606528 8.071016656 0 2.094738327 10.9160954 0 0.408222371 0.207248343

Output Stream 12a ΔH,  Enthalpy  (j/mol) 53583.913 46892.33 46698.873 0 13233.507 16244.64 0 11639.901 11289.699

Q=  nΔH 6793354.275 5146932.752 376907.3818 0 27720.73431 177328.0401 0 4751.667981 2339.771414 12529334.62

Table 2.1.57: Energy Balance For Output Isomerization Stage in FBR COMPONENT

n, MOLE RATE

Heat Vaporization, Hv

Q=  nΔH

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

126.7797347 109.7606528 8.071016656 0 2.094738327 10.9160954 0 0.408222371 0.207248343

22.305 21.916 21.916 0 40.706 15.326 0.964 6.82 5.577

2827.821982 2405.514467 176.884401 0 85.26841834 167.3000781 0 2.784085702 1.155824494 5666.729257

Table 2.1.58: Heat Vaporization at Isomerization Stage in FBR Q For Isomerization Stage in FBR = nΔH2 + nΔHv - nΔH1 = (H12a+HV)-H12 =[(12529334.62+5666.729257)-1129962.79] = 11405038.56 kJ/hr

96 2.4.9 Mass Balance in Distillation Column

D-103 Stream 12a 14 15

Temperature(0C) 400 30 34

Description Inlet Outlet Outlet

Pressure(atm) 3.0 4.5 4.0

Table2.1.59: The Operation Condition of Distillation Column

Stream 15

Stream 12a

Stream 14

Function To separate 1-butene as a byproduct and recycle n-butane back to the catalytic dehydrogenation column Condition Phase

: Liquid phase

Operating Temperature

: 300

Operating Pressure

: 3.5 atm

o

C

97

Mass Flow rate(kg/hr) Component C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Inlet S12a 7368.818521 6158.340946 452.8405315 0 37.73671096 480.4173588 0 13.06270764 5.80564784 ∑14517.02242

Outlet S15 S14 6899.898631 404.8301001 0 6190.556591 461.7197582 0 0 0 37.73671101 0 0 500.1679444 0 0 0 13.51805255 0 4.98033515 ∑14517.02242

Table 2.1.60: Inlet and Outlet Streams

2.4.10 Energy balance for Distillation Column

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 126.7797347 109.7606528 8.071016656 0 2.094738327 10.9160954 0 0.408222371 0.207248343

Input Stream 12a ΔH,  Enthalpy  (j/mol) 53583.913 46892.33 46698.873 0 13233.507 16244.64 0 11639.901 11289.699

Q=  nΔH 6793354.275 5146932.752 376907.3818 0 27720.73431 177328.0401 0 4751.667981 2339.771414 12529334.62

Table 2.1.61: Energy balance for Input Distillation Column

98

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 118.7120181 0 8.229271895 0 2.09473833 0 0 0 0

Output Stream 14 ΔH,  Enthalpy  (j/mol) 490.207 0 441.451 0 168.254 0 0 0 0

Q=  nΔH 58193.46228 0 3632.820308 0 352.4481029 0 0 0 0 62178.73069

Table 2.1.62: Energy Balance for Output Distillation Column

COMPONENT C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

Mole Rate 6.965058584 110.3348351 0 0 0 11.36487036 0 0.422452344 0.177786569

Output Stream 15 ΔH,  Enthalpy  (j/mol) 887.404 780.367 0 0 0 358.833 0 270.038 262.032

Q=  nΔH 6180.820847 86101.66424 0 0 0 4078.090524 0 114.078186 46.58577018 96521.23957

Table 2.1.63 Energy Balance for Output Distillation Column COMPONENT

n, MOLE RATE

Heat Vaporization, Hv

Q=  nΔH

C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2 TOTAL

118.7120181 0 8.229271895 0 2.09473833 0 0 0 0

22.305 21.916 21.916 0 40.706 15.326 0.964 6.82 5.577

2647.871564 0 180.7148108 0 85.26841846 0 0 0 0 2913.854793

Table 2.1.64: Heat Vaporization at Distillation Column

99 Q For Distillation Column = nΔH2 + nΔHv - nΔH1 = (H14+H15 +HV)-H12a = [(62178.73069+96521.23957+2913.854793)12529334.62] = -12367720.8 kJ/hr

Component

Amount of Energy (kJ/hr)

Fixed Bed Reactor

-127740.27

Condenser (air cooled)

1696101291

Extractive Distillation Column

697354.8662

Distillation Column

-962682.24

Compressor

14367279.35

Heat Exchanger

-41572784.51

Absorption Column

-14015183.09

Table 2.1.65: Overall Energy Balance

100

CHAPTER 3

UTILITIES

3.0

Introductions Plant utilities are a vital service to the operation of most major industrial plants. In the chemical industries, power is supplied primarily in the form of electrical energy. Agitators, pumps, hoists, blowers, compressors and similar equipment are usually operated by electric motors, although other prime movers such as steam engines, internal combustion engines, and hydraulic turbines are sometimes employed (Peters et. al, 2004).

3.1

Utilities in Plant

3.1.1 Electricity Power can be transmitted into various forms, such as mechanical energy, electrical energy, heat energy and pressure energy. In our plant, electrical is supplied by Paka Power Plant since it is nearest to our plant. The table 3.1 below shows the consumption of electricity used by our plant.

101 Equipment

Power, kW

C-101A/B

3.2424

C-102A/B

25.7246

C-103A/B

17.0327

C-104A/B

93.6073

C-105 A/B

20.8281

C-106A/B

26.03

C-107 A/B

3908.5

P-101 A/B

3.96

P-102 A/B

2.1303

P-103A/B

0.4235

P-104 A/B

49.56

P-105 A/B

0.697

P-106 A/B

0.697

P-107A/B

0.362

E-101

4842

E-102

4842

Table 3.1: Summary of consumption of electricity used by equipments

3.1.2 Water Water for industrial purpose can be obtained from one of two general sources:  the  plant’s  own  source  or  municipal  supply.  In  our  plant,  we  used  water   supply from Terengganu water company (SATU). The table 3.2 below shows the consumption of water used for operation in plant.

102 Equipment

Volume of water,m3

R-101 E-103

5.0314

E-104

7.8194

E-105

4.67

E-106

2.06

E-107

2.06 Table 3.2: Summary of consumption of water need by equipments

3.1.3 Steam Steam is supplied at unit boiler or reboiler to heat the hydrocarbon in the distillation system in our plant. The table 3.3 below shows the consumption of steam used in the plant.

Equipment

Mass of steam, kg/hr

E-101

4 514.0729

E-102

54 603.56 Table 3.3 Summary of consumption of steam used in plant

3.1.4 Air

Air is supplied to catalytic dehydrogenation of n-butane and oxidative dehydrogenation of 1-butene process. The amount of air used is 12 407.754 kg/hr.

103 3.1.5 Cooling tower

Table 3.4 below shows the amount of cooling water needed for the heat exchangers.

Equipment E-101 E-102 E-103 E-105 E-106 TOTAL

Cooling water mass flow rate (kg/hr) 25540.2 15300.4 5022.363 15057.033 2062.809 62982.8

Cooling water volume flowrate (m3/hr) 25.54 15.3 5.02 15.06 2.06 62.98

Table 3.4: Summary of consumption of cooling water need by equipments Cooling towers are a very important part of many chemical plants. They represent a relatively inexpensive and dependable means of removing low grade heat from cooling water.

Figure 3.1: Closed Loop Cooling Tower System in Butadiene production plant

104 The make-up water source is used to replenish water lost to evaporation. Hot water from heat exchangers is sent to the cooling tower. The water exits the cooling tower and is sent back to the exchangers or to other units for further cooling.

Types of Cooling Towers Cooling towers fall into two main sub-divisions: natural draft and mechanical draft. Natural draft designs use very large concrete chimneys to introduce air through the media. Due to the tremendous size of these towers (500 ft high and 400 ft in diameter at the base) they are generally used for water flow rates above 200,000gal/min. Usually these types of towers are only used by utility power stations in the United States. Mechanical draft cooling towers are much more widely used. These towers utilize large fans to force air through circulated water.

The water falls downward over fill surfaces which help

increase the contact time between the water and the air. This helps maximize heat transfer between the two. Types of Mechanical Draft Towers

Figure 3.2: Mechanical Draft Counterflow Tower

Figure 3.3: Mechanical Draft Crossflow Tower

105 Mechanical draft towers offer control of cooling rates in their fan diameter and speed of operation. These towers often contain several areas (each with their own fan) called cells.

Assumption  Cooling Tower is in cylindrical shape.  Ratio of L:D = 3:1  Volume of cooling tower, V=20m3

Calculation V=20m3 πD2/4  ×L=20 πD2/  4  ×  3D  =  20   (3D=L) D= 2.91m & L= 8.74m

106 AIR OUTLET

FAN

L

=8.74m

P-7INLET WATER

FILL AREA

AIR INLET AIR WATER

WATER OUTLET

D=2.91m Fig.3.4: Preliminary design of cooling tower

AIR INLET

107

CHAPTER 4

EQUIPMENT SIZING

4.1 DESIGN OF REACTOR

4.1.1 Multitubular Fixed Bed Reactor (R-101)

Characteristic of catalyst and bed: Catalyst used: Chromia alumina ρc = density of catalyst particle = 5000 kg/m3 Φ  =  bed  porocity  =  assumed  0.5 ρb = bulk density of bed  (ρc[1-Φ])  =  2500  kg/m3 Sizing Calculation: a) Total volume inside the tubes, Vt i)

Calculate density of mixture

Component, x

ρx (kg/m3),

Mass flow rate, (kg/hr), mx

C4H10

4.2

6885.3731

H20

0.596

5341.7169

O2

1.308

4303.5621

N2

1.251

8625.2254

H2

0.0899

2155.5521

108 ρmix =    Σ  (ρx x mx) mtotal ρmix = (4.2 x 6885.3731) + (0.596 x 5341.7169) + (1.308 x 4303.5621) + (1.251x8625.2254) + (0.0899 x 2155.5521) 37 018.135 ρmix = 1.3160 kg/m3

ii)

Calculate volumetric flow rate, vo

vo = mass flow rate density of mixture vo =

37 018.135 kg/hr 1.3160 kg/m3

vo = 28 129.2819 m3/hr

iii)

From the handbook of Elements of Chemical Reaction Engineering 4th edition, the V exact volume of reaction fluid (not including catalyst surface) :  Refer to patent US 7 488 857 B2, GHSV is ranging between 500/hr-200/hr So, we decide to choose the best value GHSV = 600/hr GHSV = Vo/V V = Vo/ GHSV = 28 129.2819 m3/hr 600/hr V = 46.8821 m3

109 iv)

By assuming that the void fraction of the catalyst bed as 50% of the total volume of reactor: Φ  =  0.5

v)

Total volume inside the tube, Vt: Φ = V/Vt Vt =  V  Φ Vt = 46.8821 m3/0.5 Vt = 93.7643 m3

b) Diameter, height and cross section area of reactor: i)

Based on engineering practice, we assume ratio of height to diameter is 1:3 So, H = 3D

ii)

Reactor is in vertical cylindrical shape: V =  πd2h/4 d = 3√4V/3π d = 3√(4  x  93.7643  m3)/  3π d = 3.4141 m

110 iii)

Calculate value of h: h = 3(3.4141 m) h = 10.2424 m

iv)

Total cross section area of reactor: Ac  =  πd2/4 Ac  =  (π  x  3.41412)/4 Ac = 9.1547 m2

c) Weight of catalyst:  Refer to the handbook of Elements of Chemical Reaction Engineering 4th edition, Wc = (1- Φ)Acz  x  ρc  Since we are more interested in catalyst weight rather than the distance z down to the reactor, we can say at the end of reactor, z = L: Wc = (1-0.5) x 9.1547 m2 x 10.2424 m x 1.3160 kg/m3 Wc = 117 207.6241 kg d) Wall thickness:  Based on table 12.10 in handbook of Plant Design & Economics for Chemical Engineering :

111  Recommended stress values (S, kPa) for SS 304 at 550 o C is 68 392.1933 kPa  Ej for spot examined if lap-welded is 0.80  Corrosion allowance (Cc) about 3 mm for a 10-year life

t=

Pri

+ Cc

Sej – 0.6 p 0.385 Sej = 0.385(68 392.1933 x 0.80) = 21 064.7955 kPa Since  that,  P  =  101.325  kPa  and  P  ≤  0.385  S ej t=

101.325 x (3.4141/2)

+ 0.003 m

(68 392.1933 x 0.80) – (0.6 x 101.325) t = 6.1648 x 10-3 m For ease the heat transfer, we will use tubes each with the following specification.  Outside diameter, do = 1 ¼ inch = 0.0318 m  Inside diameter, di = 1 inch = 0.0254 m  Maximum height of tube is 6 m

Maximum weight of catalyst fills able into a single tube: Wm =  (π  x  0.02542)/4 m2 x 6m x 5000 kg/m3 = 15.2012 kg/tube No of tube required = 117 207.6241 kg/ 15.2012 kg/tube = 7710.4068 tubes ≈  7710  tubes

112 Reactor heat transfer coefficient:

a) Tube inside heat transfer coefficient : Condition of tube region:  T = 550OC  P = 1 atm Component, x

Mass

flow Viscocity,   μx Thermal

rate,

mx (kg/s.m)

conductivity,

(kg/hr)

Heat capacity, Cpx(kJ/kg.K)

kx (W/m.K)

C4H10

6885.3731

6.82 x 10

H20

5341.7169

O2

-8

0.0160

0.096

1.295 x 10-5

0.025

1.888

4303.5621

0.0206 x 10-3

0.0266

0.9192

N2

8625.2254

0.0177 x 10-3

0.0258

1.038

H2

2155.5521

8.9060 x 10-3

0.1790

14.26

i)

Fluid  velocity  at  bulk  temperature,  μ: μ  =  Σ  (μx x mx)/mtotal μ  =  5.2699 x 10-4 kg/s.m

ii)

Fluid thermal conductivity, k: k  =  Σ  (kx x mx)/mtotal k = 0.0261 W/m.K

iii)

Heat capacity, Cp: Cp =  Σ  (Cpx x mx)/mtota Cp = 1.4694 kJ/kg.K

113 iv)

Mass velocity for each tube, Gt (kg/m2.s) Gt = mass flow rate in tubes/area inside per tube Gt  =  37  018.135  kg/hr  /  (π  x  0.02542/4 x 7710 tubes) x 1/3600 s Gt = 2.6321 kg/m2.s

v)

Pranditl number, Pr = Cpμ/k Pr = (1.4694 kJ/kg.K x 5.2699 x 10-4 kg/s.m) / 0.0261 W/m.K Pr = 29.6689

Reynold number = Gtdi/μ NRe = (2.6321 kg/m2.s x 0.0254 m) / 5.2699 x 10-4 kg/s.m NRe = 126.8626 Refer to Couldson   &   Richardson’s   Chemical   Engineering   Design   Vol.   6   Handbook: Since NRe < 2100, the flow is laminar. Nu = 1.86 (RePr)0.33 (d/l) 0.33(μ/μw)0.14 Assume  that  μ  =  μw, Nu = 1.86 x (126.8626 x 29.6689)0.33 x (0.0254/6)0.33 = 4.6374

viii)

Tube side heat transfer coefficient: hi = Nuk/d hi = (4.6374 x 0.0261 W/m.K) / 0.0254 = 4.7652 W/m2.K

b) Tube side pressure drop: i)

Refer to figure 12.4 in Couldson   &   Richardson’s   Chemical Engineering Design Vol. 6 Handbook: jf = 0.0055

114 ii)

Fluid velocity: μt =  Gt/ρ  =  2.6321  kg/m2.s /1.3160 kg/m3 μt = 2.0000 m/s

iii)

Assume that tube side velocity is 2-passes

iv)

Pressure drop: ΔPdrop  =  Np  [8  jf (l/d)  (μ/μw)-m +  2.5]  x  ρμt2/2 Assume  that  μ  =  μw, Refer to Couldson   &  Richardson’s   Chemical  Engineering  Design   Vol. 6 Handbook, m for laminar flow = 0.25 ΔPdrop   =   2[2(0.0055)   (6/0.0254)   +   2.5]   x   (1.3160   kg/m3 x 22 m2/s2)/2 ΔPdrop  =  67.8724  N/m2 = 6.6985 x 10-4 atm Since the value is very small, it can be neglected

115 c) 2- shell side heat transfer coefficient i)

By assuming 2 passes and triangular pitch: Pt = 1.25do= 1.25 x 0.0318 = 0.03975

ii)

Refer to Couldson   &   Richardson’s   Chemical   Engineering   Design   Vol. 6 Handbook, from table 12.4 fore 2 passes:  k1 = 0.249  n1 = 2.207

iii)

bundle diameter, DB: DB = dO (Nt/ k1)1/ n1 = 0.0318 x (7710/0.249)1/2.207 = 3.4455 m

Jacketed Reactor

Because of this reaction is isothermal, jacket is needed to make sure no temperature change in this reactor. Water is chosen as the cooling medium.

Figure 4.1: Jacketed reactor

116

i) Jacketed Reactor Diameter Assume diameter of jacketed reactor is 20 cm more than diameter of reactor.

ii) Height of Jacketed Assume jacket only cover

of height of reactor.

iii) Jacketed Reactor Volume Assume jacketed reactor also in a cylindrical type.

. Therefore, the reactor volume is calculated as:

iv) Volume of Water in Jacketed Reactor In order to make sure temperature no change in reactor, water as a cooling

medium must be located in between jacket and reactor.

Quantity of water

needed m3

is as follow:

117 m3 4.1.2 Multitubular Fixed Bed Reactor (R-102)

Heat release of the reactor = 10×103 kW Cp (average) = 0.0433848 kJ/molK 10×103 kJ/s = 0.0433848 kJ/molK x 10371.09 x103/3600s x (Tad – 450) Tad = 618.0 °C If all heat release goes to heating the reaction products, the temperature of the outlet fluid would be 618.0 °C which is way too high from the desired 450°C. To keep isothermal operation, we must release the heat. Simple packed does not provide efficient heat removal. A multi tube packed bed reactor is a better choice in light of heat removal. A multi tube packed bed reactor is a reactor shell consisting of many parallel tubes in which the catalyst is packed. Reactant gases flows from one end into the shell, then into the tubes, get reacted; flow out from the tubes, and the reactor shell.

Reactor design

(A) General aspects Component

Molecular

Mol flow rate (kmol/hr)

weight input

Output

C4H10

58.12

246.3758717

172.9539919

C4H8-1

56.10

84.11562543

0

C4H8-2

56.10

180.5810164

81.53266795

C4H6

54.09

18.0764008

115.8681384

H2O

18.016

616.6877427

79.39939533

CO2

44.01

11.37091046

29.38943013

H2

2.016

8573.518106

6005.393937

O2

32.00

0

118.8850405

N2

28.02

640.3678968

1524.369831

118

Flow rate in, F = 10371.09 kmol/h Gas hourly space velocity, τ = 600 h-1(298K and 1 atm) Superficial velocity, v = 1m/s Step by step approach [Reference  :  Martyn  S.  Ray  and  Martin  G  Sneesby,  ‘Chemicl  engineering  Design   Project-A   case   study   approach’,   2nd   edition,   Gordon   and   Breach   Science   publishers, 1998.] 1)

Density

of

inlet

gas

(298K

and

1

atm),

= 40.9 mol/m3= 0.0409kmol/m

=

2) Volumetric flow rate (298k and 1atm), Vs=

=2.5363 ×10⁵m /h

3) Volume of catalyst needed, VB =

= 422.72 m3

4) Density of feed gas at operating conditions, = 16.86mol/m3 = 16.86 mol/m3 x 0.058123kg/mol = 0.9798 kg/m3 5) Volume of gas at operating conditions, Vg =

= 6.1513 ×10 m3/h

6) Weight of catalyst required, WB = VBρB=422.72 m3 x 1440kg/m3 = 608716.8 kg It is desired to use multi tube packed bed rector.

119 The maximum height of tube is 6m. 7) For ease of heat transfer, we will use tubes each with the following specification. Outside diameter, Do=0.0318 m, Inside diameter, Di=0.0254m, H=6m Maximum weight of catalyst fills able into a single tube, × 6m ×1440

= 4.378 kg/ tube

Number of tubes required, Nt =

=

= 139040 tubes

To keep the reactor diameter less than 6m, we use 4 parallel reactors, each will take = 34760 tube per reactor. For a single reactor, the packing efficiency can be assumed to be 99% due to many tubes. The tubes are to be arranged in a square pitch with a tube pitch (tube centre to tube centre) of 0.0397m. 0.99πDs2/4 m2= 0.0397m2 x 34760 tubes. The reactor shell diameter, Ds = 6.49m. We will allow 1m above and below the tube bundle. Therefore the reactor will be 8m high and with internal diameter 6.49m.

The average values of the properties of fluid in the reactor: inlet

Outlet

Average

0.04365772

0.04311188

0.0433848

Viscosity, μ(Pas)

2.9452e-5

2.912e-5

2.9286e-5

Density,  ρ(kg/m3)

694.9

682.3

688.6

Heat capacity, Cp(kJ/molK)

120

Thermal conductivity of air k air = 0.05644W/mK Thermal conductivity of steam, k steam = 0.05705W/mK

Inlet molar ratio:

=

Outlet molar ratio:

= 0.9373

=

= 0.9898

Steam and air are the major components in feed and product gas, take their individual thermal conductivities to find an average conductivity. Thermal conductivity of feed, kin = 0.9373 x 0.05644 + 0.0627 x 0.05705 = 0.05648 W/mK Thermal conductivity of outlet gas, kout = 0.9898 x 0.05644 + 0.0102 x 0.05705 = 0.05645 W/mK Prantl number: Prair = 0.688; Pr steam= 1.0075 Bed  void  fraction,  ɛ  =  0.40 Mass flux of feed, G =

=

= 0.3035

8) Tube side pressure drop, Dp = catalyst pellet diameter = 6.0mm

)(

)

)( )

= 17.74Pa/m

121 L = 6m ΔPt  =  0.1064  kPa  which  is  negligible 9) Tube side heat transfer coefficient

Kreith and Bohn(1986) provide a correlation for tube side flow based on Nusselt Number Nu = 0.203(Re×Pr) 0.33+0.22Re 0.8 Pr 0.4 = 0.203(263.23 x 0.02251)0.33+0.22(263.23 0.8 x 0.022510.4) = 4.5304 hi = Nuk/D = 4.5304 x 0.05645/0.0254 = 10.07 W/m2K

Shell side properties Heat transfer salt Hitec is used to cool the tubes from the shell side. Hitec inlet temperature: 144°C Hitec outlet temperature: 600°C Tube pitch = 0.0397m P = baffle pitch = 1 m (Five 25% baffles are used in the 6m tube height) Cp = 1.3395 kJ/kg°C

μ(600°C)  =  1.2x10-3 Pas μ(144°C) = 20x10-3 Pas μ(average) =10.6 x10-3 Pas Nb = the number of tubes in the baffle window Fb= fraction of cross sectional area of shell occupied by baffle window Ds = inside diameter of reactor shell

122 Do=outside diameter of tube 10)

Shell side heat transfer coefficient Nb = fb x Nt=0.1955x34760 tubes

-

-

=1.0701 m2

Sc = PDs(1-Do/p)=1x6.49(1-0.0318/0.0397)=1.2915m2 4.9958kg/ms2

Ge  =  √(Gb Gc )=  √(4.1394 ×4.9958) =4.5475 kg / (m2 s) µ (144° C) =20×10⁻3 Pas ; µ w at wall = µ (500° C) =1.2×10-3 Pas McCabe and Smith 2005 gives 0.6

0.33

0.6

0.33

)0.14 0.14

0.6

0.33

0.14

= 39.848 W/m2K

11)

Shell side pressure drop De = effective outside diameter of tube=0.99Do=0.99x0.0318=0.03148 m Gs =mass flux of salt N = number of baffles

ρs=  average  salt  density  = Gs = 5.446kg/s ÷ (0.01x6.49mx1m) =83.91 kg/m2 s

m

123 = 242.31Pa

= 12) Overall heat transfer coefficient

ks = thermal conductivity of 1 % chromium steel r = radius of tube Heat transfer coefficient through the wall,

= 2.183× 10⁴

=

Assume no fouling,

U

8.04 W/ (m2K)

(B) Reactor shell design (1)Operating and design conditions Temperature : Normal = 450oC Peak = 500oC Tensile strength of chromium steel = 500MPa Pressure: Normal = 2atm Peak = 3atm Design pressure = (2+18 atm)x1.1 = 22 atm [18 atm is the maximum internal explosion pressure expected] Joint efficiency : Longitudinal joints and end joints : 90% Circumferential joints : 90%

124

Wall thickness

= 0.01611m = 16.11mm

P = design pressure = 22atm t = minimum wall thickness required Ds = shell inside diameter F = tensile strength of wall material(chromium steel) = 500MPa η = joint efficiency = 90% Corrosion allowance for Hitec salt is 0.2mm/year for 20 years and 25% safety factor are considered. Therefore the actual minimum thickness required for reactor wall = 16.11mm +20(0.2) +0.25x16.11mm =24.14≈25 mm

(2)Ends Torispherical ends are not suitable at the required working pressure. Either ellipsoidal or spherical ends could be used but ellipsoidal ends are preferred to minimize the overall height of the reactor and material cost. The minimum thickness for 2.5:1 ellipsoidal ends is given by

= 22.04mm

K=1.37 for 2.5 : 1 ends, P = design pressure = 22 atm Use a safety factor of 20%, t = 1.2x17.52 = 21 mm

125

The MAWP for the ellipsoidal ends is

=2.2321×10⁶ Pa = 22.03atm

The MAWP for ellipsoidal ends coincide with the cylindrical section of the reactor.

(3) Tube plate design

A tube plate is required to hold the multi parallel tubes in place at the bottom of the tube bundle. The method recommended in the British Standard BS5500 uses: CD t p

=  0.45×6.49×√

= 0.2023m= 20.23mm

Incorporate 30% safety factor, the installed tube plate will be 25 mm thick. The tube plate can be supported by 6 right angled triangular plates, equally spaced at the circumference of the reactor wall. The non-hypotenuse sides of the triangular plate are each welded to the internal shell wall and to the the bottom of the tube plate. The detailed dimensions and joining methods are subject to detailed mechanical analysis.(eg force sustained, etc)

(4) Reactor openings

The reactor openings are pipes connected to the reactor for fluid flow. The Harker equation (1978),

[(76982.11) ^0.45]

= 1492mm W = mass flow rate(kg/h) Dopt = optimum pipe diameter (mm)

126 ρ = average fluid density in the inlet and outlet pipes(kg/m3) = 0.6886 kg/m3 Insulation

Insulation is selected based on the type and size of vessel, the operating conditions, safety considerations, and insulation cost. Without insulation, the heat loss from the cylindrical section is Q  =  UAΔT =  25(π  DsL)(ΔT) =  25(π  x  6.49x6)(  0.5(144 + 600) -30) =1062kW Q = overall heat transfer coefficient for ambient conditions = 25W/m2K A= heat transfer area= inner circumferential area of shell ΔT = temperature difference between surrounding and inside the reactor shell The inside shell temperature is taken to be the average salt temperature =372oC The ambient temperature is 30oC. As calculated, the heat loss is a considerable amount of energy. By insulation, this heat loss can be integrated into the process line. Mineral wool blankets are typically used for large diameter vessel. They are as effective as they are economical. We will use two blankets of mineral wool, each 75mm thick to cover both the walls and the ellipsoidal ends.

127

End Thickness

Air Inlet

= 21mm Gas Outlet Wall Thickness = 20mm

Salt Inlet 34760Tubes I.D= 25.4mm O.D= 31.8mm Salt Outlet

Gas Inlet

Tube Plate (25mm) And Supports

128

4.1.3 Fixed Bed Reactor for Isomerazation process (R-103)

Characteristic of catalyst and bed:Catalyst used: Zeolite ρc = density of catalyst particle = 2.32 kg/m3 Φ  =  bed  porosity  =  assumed  0.5 ρb =  bulk  density  of  bed  (ρc[1-Φ])  =  1.16  kg/m3 The isomerazation of 2-butene Sizing Calculation: e) Total volume inside the tubes, Vt vi)

Calculate density of mixture

Component, x

Mole fraction

ρx @ 200oC and Mass flow rate, 3 atm, (kg/m3), (kg/hr), mx

C4H10

0.0536

4.2

821.2040

C4H8-2

0.8777

2.07

13447.215

CO2

0.0662

1.420

1014.2481

O2

0.0018

1.308

27.5777

N2

0.0007

1.251

10.72467

ρmix =    Σ  (ρx x mx) mtotal ρmix = (4.2 x 821.2040) + (2.07 x 13447.215) + (1.42 x 1014.2481) + (1.308x27.5777) + (1.251 x 10.72467) 15346.9695 ρmix = 2.00 kg/m3

129

vii)

Calculate volumetric flow rate, vo vo = mass flow rate density of mixture vo =

15346.9695kg/hr 2.14 kg/m3 vo = 7673.4848 m3/hr

viii)

From the handbook of Elements of Chemical Reaction Engineering 4th edition, the V exact volume of reaction fluid (not including catalyst surface) :  Refer to patent US 7 488 857 B2, GHSV is ranging between 500/hr-200/hr So, we decide to choose the best value GHSV = 600/hr GHSV=Vo/V V= Vo/ GHSV = 7673.4848 m3/hr 200/hr V= 38.36 m3

ix)

By assuming that the void fraction of the catalyst bed as 50% of the total volume of reactor: Φ  =  0.5

x)

Total volume inside the tube, Vt: Φ = V/Vt Vt =  V  Φ Vt = 38.36 m3/0.5 Vt = 76.73 m3

130

f) Diameter, height and cross section area of reactor:

v)

Based on engineering practice, we assume ratio of height to diameter is 1:3 So, H = 3D

vi)

Reactor is in vertical cylindrical shape: V = πd2h/4 d = 3√4V/3π d = 3√(4  x  76.73  m3)/  3π d = 3.19 m

vii)

Calculate value of h: h = 3(3.19 m) h = 9.58 m

viii)

Total cross section area of reactor: Ac  =  πd2/4 Ac  =  (π  x  3.192)/4 Ac = 7.992 m2

g) Weight of catalyst:  Refer to the handbook of Elements of Chemical Reaction Engineering 4th edition, Wc = (1- Φ)Acz  x  ρc

131  Since we are more interested in catalyst weight rather than the distance z down to the reactor, we can say at the end of reactor, z = L: Wc = (1-0.5) x 7.992 m2 x 9.58 m x 2.00 kg/m3 Wc = 7656336 kg h) Wall thickness:  Based on table 12.10 in handbook of Plant Design & Economics for Chemical Engineering :  Recommended stress values (S, kPa) for SS 304 at 350 o C is 76 800 kPa  Ej for spot examined if lap-welded is 0.80  Corrosion allowance (Cc) about 3 mm for a 10-year life

t=

Pri Sej – 0.6 p

+ Cc

0.385 Sej = 0.385(76 800x 0.80) Sej = 61 440 kPa Since that, P =  303.975  kPa  and  P  ≤  0.385  Sej t=

303.975 x (3.4141/2) (61 440 ) – (0.6 x 303.975)

+ 0.003 m

t = 0.0114 m For ease the heat transfer, we will use tubes each with the following specification.  Outside diameter, do = 1 ¼ inch = 0.0318 m  Inside diameter, di = 1 inch = 0.0254 m  Maximum height of tube is 6 m

132 Maximum weight of catalyst fills able into a single tube: Wm =  (π  x  0.02542)/4 m2 x 6m x 2320 kg/m3 = 7.053 kg/tube No of tube required = 7656.336 kg/ 7.053 kg/tube = 1085.543 tubes ≈  1085  tubes Reactor heat transfer coefficient: d) Tube inside heat transfer coefficient : Condition of tube region:  T = 350OC  P = 3 atm Component, x

Thermal conductivity, kx (W/m.K) 0.05169

Heat capacity, Cpx(kJ/kg.K)

C4H10

Mass flow Viscocity,   μx rate, mx (kg/s.m)E-4 (kg/hr) 821.2040 5.852

C4H8-2

13447.215

5.323

0.04774

4.46

CO2

1014.2481

6.41

0.03721

0.98651

O2

27.5777

2.977

0.04709

1.01

N2

10.72467

1.582

0.04254

1.06

vi)

Fluid  velocity  at  bulk  temperature,  μ: μ  =  Σ  (μx x mx)/mtotal μ  =  5.026E-4 kg/s.m

vii)

Fluid thermal conductivity, k: k  =  Σ  (kx x mx)/mtotal k = 0.04717 W/m.K

viii)

Heat capacity, Cp: Cp =  Σ  (Cpx x mx)/mtota Cp = 3.9808 kJ/kg.K

0.096

133 ix)

Mass velocity for each tube, Gt (kg/m2.s) Gt = mass flow rate in tubes/area inside per tube Gt  =  15346.9695kg/hr  /  (π  x  0.02542/4 x 1085 tubes) x 1/3600 s Gt = 7.7541 kg/m2.s

x)

Pranditl number, Pr = Cpμ/k Pr = (3.9808 kJ/kg.K x 5.026E-4 kg/s.m) / 0.04717 W/m.K Pr = 42.416

Reynold number = Gtdi/μ NRe = (7.7541 kg/m2.s x 0.0254 m) / 5.026E-4 kg/s.m NRe = 357.9 Refer to Couldson   &   Richardson’s   Chemical   Engineering Design Vol. 6 Handbook: Since NRe < 2100, the flow is laminar. Nu = 1.86 (RePr)0.33 (d/l) 0.33(μ/μw)0.14 Assume  that  μ  =  μw, Nu = 1.86 x (357.9 x 42.416)0.33 x (0.0254/6)0.33 = 3.9503 viii)

Tube side heat transfer coefficient: hi = Nuk/d hi = (3.9503x 0.04717 W/m.K) / 0.0254 = 7.336 W/m2.K

e) Tube side pressure drop:

v)

Refer to figure 12.24, pages 667 in Couldson   &   Richardson’s   Chemical Engineering Design Vol. 6 Handbook: jf = 0.023

vi)

Fluid velocity: μt =  Gt/ρ  =  7.7541  kg/m2.s /2.0 kg/m3 μt = 3.877 m/s

134 vii)

Assume that tube side velocity is 1-passes

viii)

Pressure drop:

ΔPdrop  =  Np  [8  jf (l/d)  (μ/μw)-m +  2.5]  x  ρμt2/2 Assume  that  μ  =  μw, Refer to Couldson   &  Richardson’s   Chemical  Engineering  Design   Vol. 6 Handbook, m for laminar flow = 0.25

ΔPdrop = 1[8(0.023) (6/0.0254) + 2.5] x (2.00 kg/m3 x 3.877 m2/s2)/2

2

ΔPdrop  =  690.9  N/m2 = 0.0069 atm Since the value is very small, it can be neglected

f) 2- shell side heat transfer coefficient

iv)

By assuming 1 passes and triangular pitch: Pt = 1.25do= 1.25 x 0.0318 = 0.03975

v)

Refer to Couldson   &   Richardson’s   Chemical   Engineering   Design   Vol. 6 Handbook, from table 12.4 for 1 passes:  k1 = 0.319  n1 = 2.142

vi)

bundle diameter, DB: DB = dO (Nt/ k1)1/ n1 = 0.0318 x (1085/0.319)1/2.142 = 1.4164 m

135 4.2

DESIGN OF MASS TRANSFER EQUIPMENT

4.2.1 Extractive Distillation (D-101)

Equilibrium stage Component Relative Feed volatility, amount at (kg/hr) (53 , 5atm)

Feed amount (kmol/hr)

Feed fraction

Distillate (di)

Bottom (bi)

C4H10 (HK) C4H8-2

0.88 1.38

10496.55 3700.44

180.975 66.0790

0.4825 0.1880

0.003 0.010

0.665 0.033

C4H6 (LK)

1.53

6578.56

121.825

0.3295

0.980

0.002

Multicomponent distillation Using Fenske Equation, Nmin =

= = 15.37 plates

16 plates

Minimum Reflux Ratio, = 1-q 

For liquid feed, q=1.0 =

136 

By using trial and error method, When

1.089

= -2.032+0.8915+1.143 = 0.0025 must lie between the 0.88

0

values of the HK and LK, HK

1.53. By trial and error,

Rmin + 1 =

= = 3.435 So, Rmin = 3.435 – 1 = 2.435 Number of equilibrium stages, R = Operational reflux ratio = 3.0

= 0.75 = 0.75 = 0.5023 N= 33.16

33 stages

With, = 9.267 kg/m = 540.3 kg/m =

= = 0.098

= 1.089

LK

137 Use a tray spacing of 0.31 m, From Figure 15.5, (Plant Design and Economic for Chemical Engineers, page 778) = 0.098, Csb = 7.50 10-1 m/s

At

The surface tension of 1,3-butadiene,

= 4.34 dyne/cm

Vnf = Csb = 7.50×10-2 = 0.4188 m/s Assume 80% of flooding, Vn = 0.8 Vnf = 0.8(0.4188) = 0.3351 m/s The net column area, An=

=

=

= 2.236 m2 Assume that the downcomer occupies 15% of the cross sectional area of the column. Thus, Area of column, Ac = = = 2.6306 m2

Diameter of column, D = = = 1.830 m

138 Number of plates = 33 – 1 = 32 plates

Functional height = (32-1) ×0.31 = 9.61 m Column Height The column height is the functional height(between top and bottom plate) plus the base and vapor space at the bottom and top of column. Assign 30% of the functional height to the column base and top mass flow.

Actual column height = 9.61 m × 1.3 = 12.493 m

4.2.2 Extractive Distillation Column D-102

Equilibrium stage Component Relative Feed volatility, amount at (kg/hr) (53 , 5atm) C4H10 (LK) 1.17 C4H81.00 2(HK) NMP -

Feed fraction

Distillate (di)

Bottom (bi)

165.8981 2.8544 14.4192 0.257

0.0107 0.00093

0.926 0.0731

0.0076 0.0095

-

-

-

-

Multicomponent distillation Using Fenske Equation, Nmin =

Feed amount (kmol/hr)

-

139

= = 17.59 plates 18 plates Minimum Reflux Ratio, = 1-q  

For liquid feed, q=1.0

=



=0

By using trial and error method, When

1.02 = -2.032+0.8915+1.143 = 0.0025

0

must lie between the values of the HK and LK, HK 1.00 1.173. By trial and error, =1.02

LK

Rmin + 1 = = = 3.345 So, Rmin = 3.435 – 1 = 2.345 Number of equilibrium stages, Method used Gilliland Correlation (‘Plant Design & economic for Chemical Engineers’,  Max  S.Peter  Klaus  D.  Timmerhaus,  Ronald,  pages 772) R = Operational reflux ratio = Rule of thumb= 1.2 (2.345)= 2.814

= 0.75 = 0.75

140

= 0.521 N

= 37.578

38 stages

Determining Actual trays Plant Efficiency Rule of thumb Tray efficiency for distillation of light hydrocarbon and aqueous solution are 60% -90% N act =

=

= 41.75

42 trays

With, = 10.455 kg/m = 591.5 kg/m Assume =

= =

= = 0.09971 Use a tray spacing of 0.31 m, From Figure 15.5, (Plant Design and Economic for Chemical Engineers, page 778)

= 0.098, Csb = 7.50 10-1 m/s

At

The surface tension of 1,3-butadiene, Vnf = Csb = 5.59×10-2 = 0.39586 m/s

= 4.34 dyne/cm

141 Assume 80% of flooding, Vn = 0.8 Vnf = 0.8(0.39586) = 0.3167 m/s The net column area, An=

=

=

= 5.1413 m2

Assume that the downcomer occupies 15% of the cross sectional area of the column. Thus, Area of column, Ac = = = 6.048 m2

Diameter of column, D = = = 2.7750 m Number of plates = 35 – 1 = 34 plates Functional height = (34-1) ×0.31 = 10.23 m

Column Height The column height is the functional height(between top and bottom plate) plus the base and vapor space at the bottom and top of column. Assign 30% of the functional height to the column base and top mass flow. Actual column height = 10.23 m × 1.3 = 13.299 m

142 4.3 DESIGN OF DISTILLATION COLUMN

4.3.1 Distillation Column D-103

Design basis:  The feed into the conventional distillation column T-102 with total mass flow rate of 15504.4987kg/hr at 116°C and 4.0 atm.  Distillation is done under 3.5 atm and 115°C at the top of the column while 4.5 atm and 144°C at the bottom of the column.

Design parameters:  Reflux ratio, R  Number of tray, N  The feed stage  Height of the column, H  Internal diameter of the column, D

Design criteria: Distillation is done under moderate pressure condition as the boiling points of C4fractions are very low.

Design calculation: 1-butene is being chosen as the light key with the overhead and n-butane is being chosen as the heavy key.

143 Determining Minimum Trays: Method  used:  Fenske’s  equation Source:  James  M.  Douglas,  “Conceptual  Desin  of  Chemical  Process”,  pg 441

Determining Minimum Reflux Ratio: Method  used:  Underwood’s  equation Source:  James  M.  Douglas,  “Conceptual  Desin  of  Chemical  Process”,  pg 442

Due to the concentration of n-butane and 1-butene is very less compared to others, thus, they are not taken into account during the calculation. As the light key enters at its boiling point, therefore, the q value is 1.

By  solving  the  equation,  Θ=  2.9390

Rule of thumb : R=1.2Rmin Determining Number of Equilibrium Trays: Method used: Gilliland Correlation Source:  Max  S.  Peter,  Klaus  D.  Timmerhaus,  Ronald  E.  West,  “Plant  Design  and   Economics for Chemical  Engineer”,  pg 772

144

N = 18.44 Determining Actual Trays-Plate Efficiency Rule of thumb:Tray efficiency for distillation of light hydrocarbons and aqueous solutions are 60-90%.

Determining the Net Vapor Velocity at Flooding Conditions and the Diameter of Distillation Column Source:  Max  S.  Peter,  Klaus  D.  Timmerhaus,  Ronald  E.  West,  “Plant  Design  and Economics  for  Chemical  Engineer”,  pg 777-778 **Note:  The net vapor velocity will be obtained at both ends of the column to determine the maximum allowable vapor velocity and hence maximum column diameter that will be required.  The properties for vapor and liquid phases in the distillation column are obtained by using ASPEN PLUS.  As the multi-component stream consists of C4 fractions, acetone and sulfolane, thus, average values of properties such as densities and surface tensions will be estimated based on the mole fraction of each component presents in the stream.

145 For top of the column :

=0.9495 Taking tray spacing as 0.91m, Csb obtained = 0.13 m/s. [From figure 15.5, “Plant  Design  and  Economics  for  Chemical  Engineer”,  pg  778]

= 0.73 m/s Assuming 80% flooding, Vn = 0.8Vf = 0.8(0.73) =0 .58

Net Area,

Assume downcomer occupies 15% of the cross sectional area of the column,

=3.24 So, the top column diameter, D=

146 It is shown that the maximum column diameter happens at the top of the column. Therefore, the diameter of column D-103 = 2.13 m. Determining the height of the column: Source:  Max  S.  Peter,  Klaus  D.  Timmerhaus,  Ronald  E.  West,  “Plant  Design  and   Economics  for  Chemical  Engineer”,  pg 779. Height of the column, = (31-1)(0.91) =27.30 m Additional height of 15% excess space is required for vapor-liquid disengaging space at the top of the column and for liquid sump at the bottom of the column. Therefore, the actual height of the column, Hc=1.15H =1.15(27.30) m =31.40 m.

147 4.4 DESIGN OF ABSORPTION TOWER

4.4.1 Absorption tower (T-101)

Stream 10

Stream 9

Stream 11

Mole Fraction Component C4H10 C4H8-2 C4H6 H2O CO2 O2 N2 H2

Inlet S9 0.1248 0.0444 0.0779 0.0091 0.0158 0.0474 0.5303 0.1503

Stream 11: 21754.4987 kg/hr Stream 9: 22219.17 kg/hr Stream 10: 464.6718 kg/hr

Outlet S10 0.0006 0.0008 0.0004 0.0029 0.0150 0.0635 0.7143 0.2025

S11 0.4825 0.1701 0.3010 0.0271 0.0179 0.0009 0.0005 0.0000

148

Component kg/hr 2772.95 986.53 1730.86 202.19 351.06 1053.19 11782.83 3339.541

C4H10 C4H8-2 C4H6 H2O CO2 O2 N2 H2

GAS STREAM Input kmol/hr 47.809 17.617 32.053 11.233 7.979 32.912 841.631 1669.77

kg/s 0.7703 0.2740 0.4808 0.05617 0.09752 0.2926 3.273 0.9277

Total: 2661.004 kmol/hr

Step 1: Selecting Type of Packing: We select the Ceramic Intalox Saddles as our packing material. The size of the packing material is 1 inch (38mm). Therefore, the packing factor Fp = 170m-1. (Adapted from Table 11.2, Coulson & Richardson’s,   ‘Chemical   Engineering   Design’  Volume  6,  page  591)

Step 2: Determine the Tower Cross-Sectional Area, AT and Diameter, DT (CORNELL’S  Method) Firstly, we assume the ratio of molar liquid flow rate to molar gas flow rate about 10. Thus, the molar flow rate required is 10 x 2661.004 = 26610.04 kmol/hr. Molar Gas Flow Rate, Vw* = 2661.004kmol/hr =6.1721kg/s Molar Liquid Flow Rate, Lw* = 26610.04kmol/hr =61.721kg/s Gas Density at 40°C,ρv = 0.03698kmol/m3=1.035kg/m3(From Aspen Properties) Liquid   Density   at   40°C,   ρL= 11.2167kmol/m3 =314.07kg/m3(From Aspen Properties) Liquid Viscosity at 40°C,μL = 3.58x10-4Ns/m2 (From Aspen Properties)

FLV =

=

= 0.574

149 By using the FLV value, from the figure 11.44 (Coulson   &   Richardson’s,   ‘Chemical  Engineering  Design’  Volume  6), we get the K4 value. We design for a pressure drop of 20mm H2O/m packing. K4 = 0.6. At flooding K4 = 1.3. From figure 11.41 and 11.42 (Coulson  &  Richardson’s,  ‘Chemical  Engineering  Design’   Volume 6, page 599 and 600)

So, Percentage of Flooding =

The Vw* value =

= 0.54kg/m2.s

The column area required, A = = 11.4298m2

 D= = = 3.815 m

Step  3:  Determine  the  Tower  Height,  H.  (CORNELL’S  Method) Kinematics Viscosity for liquid, DL = 3.19x10-7 m2/s Kinematics Viscosity for gas, DV = 4.97x10-6m2/s

150 Viscosity for gas,μv = 1.839 x 10-5 Ns/m2 = 1.839x10-5/1.035(4.97x10-6)

Gas Schmidt Number , (Sc)v =

= 3.58

= 3.58x10-4/314.07(3.19x10-7)

Liquid Schmidt Number, (Sc)L =

= 3.573

Liquid mass flow rate per unit area column cross sectional area, kg/m2.s From figure 11.43, Φn = 0.1. ( Coulson & Richardson, Chemical Engineering Design, Vol 6) For 68% flooding, from Figure 11.41 and 11.42, K3 =  0.8  and  ψn = 50 HOG can be expected to be around 1m, so we estimate of Z=20 m . For gas, f1 = f2 = f3 = 1.

Height of Liquid Phase Transfer Unit, HL = HL = 0.305 n(Sc)L . K3

= 0.305(0.1)(3.573)0.5 . 0.8 = 0.06115 m

Height of a Gas Phase Transfer Unit, HG =

HG= 0.011 n (Sc)v =

 HOG = HG+m

HL

= 2.40568 + 0.8(0.06115) = 1.331725 m

=

= 2.40568 m

151  Height of Column = NOG x HOG

estimated

= 8 x 2.4546 = 19.637 m  This value is close enough to the estimated value

× 100% =

× 100% = 1.815%

Where: K3 = percentage flooding correction factor Φn = HL factor from figure 11.43 Dc = Column Diameter, m Z = Column Height, m Ψn = HG factor from figure 11.42

f1 = Liquid viscosity correction factor = f2 = Liquid density correction factor = f3 = Surface tension correction factor =

 So, the packed bed height is 19.367 m and diameter of column is 3.815 m.

152 4.5 DESIGN OF HEAT EXCHANGER

4.5.1 Heat Exchanger (E-101)

Stream B T=100°C P=2 atm ṁ= 26540.2kg/hr

Stream 7 T=450°C P=2 atm ṁ=65140kg/hr

Stream 7a T=200°C P=2 atm ṁ=65140kg/hr

Stream A T=20°C P=2 atm ṁ=26540.2 kg/hr

 Hot fluid – mixture of reaction stream- 7  Cold fluid – water  Q = McP Δt  Mass flow of liquid mixture, m = 18.09 kg/sec  Specific heat of liquid mixture, Cp = 0.5399 KJ/Kg.K  Outlet temperature = 200oC  Inlet temperature = 450oC  Total amount of heat to be removed, Q = 2441.7 KJ/sec  Let inlet temperature of water = 20oC  Outlet temperature = 100oC

153 Specific heat of water = 4.14 KJ/Kg.K Mass flow rate of water = 2441.7 / (4.14X80) = 7.372Kg/sec @ 26540.2 Kg/hr

Routing of fluids : Water which has the high flow rate is taken in tube side. Liquid mixture which has viscosity higher than water is taken in shell side. LMTD: Liquid mixture

Water

Δt

Thi = 450 oC

Tco = 100 oC

350 oC

Tho = 200 oC

Tci = 20 oC

180 oC

LMTD = {[(Thi – Tco) – (Tho – Tci) ] / ln [(Thi – Tco) – (Tho – Tci) ] } LMTD = (350– 180)/ln(350/180) = 255.65 oC R = (Thi – Tho ) / (Tco – Tci) = (450– 200) / (100 – 20) = 3.125 S = (Tco – Tci ) / (Thi – Tci) = ( 100 - 20) / (450 – 20) = 0.186 For R = 3.125and S = 0.186 From figure 12.20, correction factor FT Coulson  and  Richardson’s  Chemical   Engineering Design Vol. 6. For two shell pass; four or multiple of four tube passes, FT = 0.92

154 (LMTD)cor = 0.92 x 255.65 = 207.6 Assumption:  (Table  12.1,’  Chem.Eng’,  Vol.6) Overall heat transfer coefficient, U = 200 W/m2 .oC (average typical coefficient from Table 12.1: Coulson  and  Richardson’s  Chemical  Engineering Design Vol. 6.) Area = Q/ (LMTD x U) = (2441.7 x 103 )/ (200 x 207.6) = 58.8 m2 Layout and tube size Using a split-ring floating head for ease of cleaning and can operate in high differential temperature. High viscosity fluid should circulate on shell side; steam has higher viscosity than steam, so put air through the tube and steam in shell. High temperature steam may cause corrosive, so choose stainless steel as the material of construction for both tubes and shells. Use 19.05 mm as outside diameter, 14.83 mm as inside diameter and 4.88 meter long tubes on a triangular 23.81 mm pitch (pitch/diameter = 1.25). (Coulson  and  Richardson’s  Chemical  Engineering  Design  Vol.  6) Calculate number of tubes Area of one tube (neglecting thickness of tube sheets) Atube =    π(outside  diameter)(length) =  π  (19.05x103)(4.88) = 0.2921 m2 So, number of tubes = A / Atube = (58.8 m2) / (0.2921 m2) = 201 tubes For four passes, tubes per pass is 50 tubes

155 Calculate bundle and shell diameter From Table 12.4: Coulson and Richardson’s  Chemical  Engineering  Design  Vol.  6   For 4 passes, KI = 0.175, nI = 2.285 So, Bundle diameter, Db = do(Nt/KI)

1/ nI

Where; do = outside diameter N t = number of tubes

Db = 19.05(201/0.175)

1/ 2.285

= 416.04mm @ 0.416m For a split-ring floating head exchanger the typical shell clearance from Figure 12.10: Coulson  and  Richardson’s  Chemical  Engineering  Design  Vol.  6  is 58 mm, so shell inside diameter, Shell diameter, Ds = 416.04mm + 58 mm = 474 mm (0.474 m)

156 4.5.2 Heat Exchanger (E-102)

Stream B T=100.8°C P=2 atm ṁ= 15300.4kg/hr

Stream 7a T=200°C P=2 atm ṁ=65140kg/hr

Stream 9 T=40°C P=2 atm ṁ=65140kg/hr

Stream B T=20°C P=2 atm ṁ=15300.4 kg/hr

 Hot fluid – mixture of reaction stream- 7a  Cold fluid – water  Q = McP Δt  Mass flow of liquid mixture, m = 18.09 kg/sec  Specific heat of liquid mixture, Cp = 0.5399 KJ/Kg.K  Outlet temperature = 200oC  Inlet temperature = 40oC  Total amount of heat to be removed, Q = 1563.07 KJ/sec  Let inlet temperature of water = 20oC  Outlet temperature = 100.8oC

Specific heat of water = 4.14 KJ/Kg.K Mass flow rate of water = 2441.7 / (4.14X80) = 4.25Kg/sec @ 15300.4 Kg/hr

157 Routing of fluids : Water which has the high flow rate is taken in tube side. Liquid mixture which has viscosity higher than water is taken in shell side. LMTD: Liquid mixture

Water

Δt

Thi = 200 oC

Tco = 100.8 oC

99.2 oC

Tho = 40 oC

Tci = 20 oC

20oC

LMTD = {[(Thi – Tco) – (Tho – Tci) ] / ln [(Thi – Tco) – (Tho – Tci) ] } LMTD = (200– 100.8) ln (200/100.8) = 144.78 oC R = (Thi – Tho ) / (Tco – Tci) = (200– 40) / (100.8 – 20) = 1.98 S = (Tco – Tci ) / (Thi – Tci) = ( 100.8 - 20) / (200 – 20) = 0.449 For R = 1.98and S = 0.449 From figure 12.20, correction factor FT Coulson  and  Richardson’s  Chemical   Engineering Design Vol. 6. For two shell pass; four or multiple of four tube passes, FT = 0.85 (LMTD)cor = 0.85 x 144.78 = 123.06 Assumption:  (Table  12.1,’  Chem.Eng’,  Vol.6)

158 Overall heat transfer coefficient, U = 200 W/m2 .oC (average typical coefficient from Table 12.1: Coulson  and  Richardson’s  Chemical  Engineering Design Vol. 6.) Area = Q/ (LMTD x U) = (1563.07 x 103 )/ (200 x 123.06) = 63.5 m2 Layout and tube size Using a split-ring floating head for ease of cleaning and can operate in high differential temperature. High viscosity fluid should circulate on shell side; steam has higher viscosity than steam, so put air through the tube and steam in shell. High temperature steam may cause corrosive, so choose stainless steel as the material of construction for both tubes and shells.

Use 19.05 mm as outside diameter, 14.83 mm as inside diameter and 4.88 meter long tubes on a triangular 23.81 mm pitch (pitch/diameter = 1.25). (Coulson  and  Richardson’s  Chemical  Engineering  Design  Vol.  6) Calculate number of tubes Area of one tube (neglecting thickness of tube sheets) Atube =    π(outside  diameter)(length) =  π  (19.05x103)(4.88) = 0.2921 m2 So, number of tubes = A / Atube = (63.5m2) / (0.2921 m2) = 217.4 tubes For four passes, tubes per pass is 54 tubes

159 Calculate bundle and shell diameter From Table 12.4: Coulson and  Richardson’s  Chemical  Engineering  Design  Vol.  6   For 4 passes, KI = 0.175, nI = 2.285 So, Bundle diameter, Db = do(Nt/KI)

1/ nI

Where; do = outside diameter N t = number of tubes

Db = 19.05(217.4 /0.175)

1/ 2.285

= 430.6mm @ 0.43m For a split-ring floating head exchanger the typical shell clearance from Figure 12.10: Coulson  and  Richardson’s  Chemical  Engineering  Design  Vol.  6  is 58 mm, so shell inside diameter, Shell diameter, Ds = 430.6 mm + 58 mm = 488 mm (0.488 m)

160 4.5.3 Heat exchanger (E-103)

Step 1: Specification

Stream 2 T=24°C P=4 atm ṁ=5022.363kg/hr

Stream 13a T=30°C P=4 atm ṁ= 6250kg/hr

Stream 13 T=40°C P=4 atm ṁ=6250 kg/hr

Stream 1 T=20°C P=4 atm ṁ=5022.363kg/hr

QS13 = QS1 =  ṁc∆T  =  83853.36747  kJ/hr                   (ṁc∆T)water = 83853.36747 kJ/hr ṁ  =   = 5022.363 kg/hr of cooling water

Step 2: Overall Coefficient For this type of heat exchanger for overall coefficient will be in the range 20-300 W/m.K which is from (Table 12.1), (Coulson & Richardson, Chemical Engineering Design, Vol 6). So, we start with average value at 150 W/m.K.

161 Step 4: Exchanger Type and Dimensions With, Th1= 40º C Th2= 30º C Tc1= 20º C Tc2= 24º C

∆Tm =

= = 12.77 ºC

R= = = 2.50

S= = = 0.20

Heat duty,Q = 83853.36747kJ/hr X 1hr/3600s = 23.293 kW From Table 12.19, (Coulson & Richardson, Chemical Engineering Design, Vol 6), at S=0.20 and R=2.5, FT=0.96 which is acceptable.

162 ∆Tm=  FT∆Tlm = 0.96(12.77) = 12.26º C

Step 5: Heat Transfer Area Heat transfer area = A0 = = = 12.666m2

Step 6: Layout and tube sizes Neither fluid is corrosive and the operating pressure is not high, so a plain carbon steel can be used for shell and tubes. The hot stream has large flowrate, so let put it in the tube side and cold stream in the shell side and then Use , Outside diameter, d0 = 19.05 mm Inside diameter, di = 16.56 mm Richardson, Engineering Design, Vol 6) Length, L = 5.00 m Tube pitch = square for easy cleaning, 1.25do = 1.25(19.05) = 23.8125 mm

Table 12.3, (Coulson & Chemical

163 Step 7: Number of Tubes Area of one tubes (neglecting thickness of tube sheets) =

d0 L

= (19.05mm)(5.0mm) = 0.2992 m2

Total number of tube required = = 42 tubes

So for 1 passes, total number of tubes required per pass = 42 tubes Check the tube side velocity at this stage: Tube cross sectional area =

= 0.0002154 m

So, the area per pass = 42×0.0002154 = 0.0090468 m



Volumetric flow rate =

×

= 0.002778 m /s 

Tube side velocity, ut = = = 0.3071 m/s

164 Step 8: Bundle and Shell diameter From Table 12.4, for 1 tube passes K1=0.215, n1=2.207 Source: Table 12.4 page 649 Chemical  Engineering  Design,  Vol  6,  Coulson  &  Richardson’s Constant for use in the heat exchanger calculation So, Tube bundle diameter,Db = do = 19.05 = 207.91 mm = 0.208 m For a fixed-U tube heat exchanger, the typical shell clearance from Figure 12.10, (Coulson & Richardson, Chemical Engineering Design, Vol 6, page 663) is 10 mm. So, the shell diameter, Shell inside diameter, Ds = (207.91 mm + 10 mm) = 217.91 mm = 0.218 m

165 4.5.4 Condenser (E-106)

Step 1: Specification

QS12b =  ṁc∆T  =942720.8361  kJ/hr   (ṁc∆T)water = 942720.8361 kJ/hr

ṁ  =   = 15057.033 kg/hr of cooling water Heat duty, Q = = 202.5773KW

x 2.38 x (60-40)

166 Step 2: Physical properties Properties

Unit

Inlet

Mean

Outlet

Hot stream Flow

kg/hr

Temperature

0

Specific heat

kJ/kg.K

Thermal

W/m.K

C

Conductivity Density

Kg/m3

Viscosity

M.Ns/m2

15320.96963 60

50

40

2.42

2.38

2.34

0.090505

0.09491

0.099315

418

504.75

591.5

0.26022

0.20317

0.14612

Cold Stream

15057.033

Flow

kg/hr

Temperature

0

Specific heat

kJ/kg.K

Thermal

W/m.K

C

Conductivity Density

Kg/m3

Viscosity

M.Ns/m2

20

27.5

35

27.5

4.185

4.18

0.1813

0.179195

4.19

4.185

4.18

1000

1000

1000

0.83547

0.74244

0.64941

Step 3: Overall Coefficient For this type of heat exchanger for overall coefficient will be in the range 350-900 W/m.K which is from (Table 12.1), (Coulson & Richardson, Chemical Engineering Design, Vol 6). So, we start with average value at 350 W/m.K.

Step 4: Exchanger Type and Dimensions With, Th1= 60º C Th2= 40º C Tc1= 35º C

167

Tc2= 20∆Tm =

= = 22.407 ºC

Then, R= = = 1.33

S= = = 0.375

From Table 12.19, (Coulson & Richardson, Chemical Engineering Design, Vol 6),pages 657 at S=0.375. and R=1.33, FT=0.90 which is acceptable.

∆Tm=  FT∆Tlm = 0.90(22.407) = 20.17º C

168 Step 5: Heat Transfer Area

Heat transfer area = A0 = = = 28.7m2 Step 6: Layout and tube sizes

Neither fluid is corrosive and the operating pressure is not high, so a plain carbon steel can be used for shell and tubes. The hot stream has large flowrate, so let put it in the tube side and cold stream in the shell side and then Use , Outside diameter, d0 = 19.05 mm Inside diameter, di = 14.83 mm

Table 12.3, (Coulson &

Richardson,

Chemical

Engineering Design, Vol 6) Length, L = 5.00 m Tube pitch = square for easy cleaning, 1.25do = 1.25(19.05) = 23.8125 mm

Step 7: Number of Tubes Area of one tubes (neglecting thickness of tube sheets) =

d0 L

= (19.05mm)(5.0mm) = 0.2992 m2

169

Total number of tube required = = 95.98 tubes ~ 96 tubes

So for 1 passes, total number of tubes required per pass = 96 tubes Check the tube side velocity at this stage: Tube cross sectional area =

= 0.00017273 m

So, the area per pass = 95×0.00017273 m = 0.016582 m



Volumetric flow rate = = 0.002778 m /s



Tube side velocity, ut = = = 0.508 m/s

×

170 Step 8: Bundle and Shell diameter From Table 12.4, for 1 tube passes K1=0.215, n1=2.207 Source: Table 12.4 page 649 Chemical  Engineering  Design,  Vol  6,  Coulson  &  Richardson’s Constant for use in the heat exchanger calculation So, Tube bundle diameter,Db = do = 19.05 = 302.3738 mm = 0.3023 m For a fixed-U tube heat exchanger, the typical shell clearance from Figure 12.10, (Coulson & Richardson, Chemical Engineering Design, Vol 6, page 663) is 10 mm. So, the shell diameter, Shell inside diameter, Ds = (0.3023 + 0.053 m) = 0.3553 m

Step 9: Tubes side heat transfer coefficient

Re =

= =18716.37 Prandt Number

Pr =

171

= = 5.0947

L/di = = 337.154 From figure 12.23 (Coulson & Richardson, Chemical Engineering Design, Vol 6, page 665) Re =18716.37 and L/di = 337.154 jh = 0.0040

Nu = jh x Re x Pr1/3 = 0.0040 x 18716.37 x (5.0947)1/3 = 128.8213

hi =

=

= 824.439

172 Step 10: Shell side heat transfer coefficient Kern’s  method  will  be  used Db = 0.3023 m

refer step 8 same equation

Ds = 0.3553 m

Baffle spacing, lb Ɩb = = = 0.07106 m Tube pitch Pt = 1.25 do = 1.25 x 0.01905 m = 0.0238125 m Cross sectional,As As =

= = 5.0495x 10-3 m2 Equivalent diamater for square pitch :de = = = 0.01881 m

173 Volumetric flowrate on shell side :V=

x

x

= 0.0043075 m3/s

Shell side velocity

=

= = 0.85305 m/s

Re =

= = 21612.3465

Prandt Number

Pr = = = 17.3393 Choose 25% battle cut From figure 12.29 (Coulson & Richardson, Chemical Engineering Design, Vol 6, page 672) Re =21612.3465

174 jh = 0.0014 Neglecting viscosity correction item hs=

x jh x Re x Pr 1/3

x 0.0014 x 21612.3465 x 17.33931/3

=

= 746.0666 m2/K

Step 11: Overall coefficient

=

=

+

+

+

+ 0.0005 +

=

= 244.8514 W/m2.K

+

+

+

175 4.5.5 Condenser (E-107)

Step 1: Specification

QS13 = QS1 =  ṁc∆T  =  83853.36747  kJ/hr (ṁc∆T)water = 83853.36747 kJ/hr

ṁ  =   = 5022.363 kg/hr of cooling water

Step 2: Overall Coefficient For this type of heat exchanger for overall coefficient will be in the range 20-300 W/m.K which is from (Table 12.1), (Coulson & Richardson, Chemical Engineering Design, Vol 6). So, we start with average value at 150 W/m.K.

Step 3: Exchanger Type and Dimensions With, Th1= 40º C Th2= 30º C Tc1= 20º C Tc2= 24º C

176

∆Tm =

=

= 12.77 ºC

R= = = 2.50

S= = = 0.20

Heat duty,Q = 83853.36747kJ/hr X 1hr/3600s = 23.293 kW

From Table 12.19, (Coulson & Richardson, Chemical Engineering Design, Vol 6), at S=0.20 and R=2.5, FT=0.96 which is acceptable.

∆Tm=  FT∆Tlm = 0.96(12.77) = 12.26º C

177 Step 5: Heat Transfer Area

Heat transfer area = A0 = = = 12.666m2

Step 6: Layout and tube sizes Neither fluid is corrosive and the operating pressure is not high, so a plain carbon steel can be used for shell and tubes. The hot stream has large flowrate, so let put it in the tube side and cold stream in the shell side and then Use , Outside diameter, d0 = 19.05 mm Inside diameter, di = 16.56 mm

Table 12.3, (Coulson &

Richardson,

Chemical

Engineering Design, Vol 6) Length, L = 5.00 m Tube pitch = square for easy cleaning, 1.25do = 1.25(19.05) = 23.8125 mm

Step 7: Number of Tubes Area of one tubes (neglecting thickness of tube sheets) =

d0 L

= (19.05mm)(5.0mm) = 0.2992 m2

178

Total number of tube required = = 42 tubes So for 1 passes, total number of tubes required per pass = 42 tubes Check the tube side velocity at this stage: Tube cross sectional area =

= 0.0002154 m

So, the area per pass = 42×0.0002154 = 0.0090468 m



Volumetric flow rate =

×

= 0.002778 m /s 

Tube side velocity, ut = = = 0.3071 m/s

Step 8: Bundle and Shell diameter From Table 12.4, for 1 tube passes K1=0.215, n1=2.207 Source: Table 12.4 page 649 Chemical  Engineering  Design,  Vol  6,  Coulson  &  Richardson’s Constant for use in the heat exchanger calculation So,

Tube bundle diameter,Db = do = 19.05 = 207.91 mm = 0.208 m

179

For a fixed-U tube heat exchanger, the typical shell clearance from Figure 12.10, (Coulson & Richardson, Chemical Engineering Design, Vol 6, page 663) is 10 mm. So, the shell diameter, Shell inside diameter, Ds = (207.91 mm + 10 mm) = 217.91 mm = 0.218 m

4.6 DESIGN OF REBOILER

4.6.1 Reboiler (E-104)

   

This is a kettle reboiler using U-tube Do = 30mm, Di = 25mm Nominal tube length = 10.4 m Inlet temperature of bottom product = 34

 Desired temperature for vapor = 110  Heating medium = Downtherm A, 160

Step 1: Specification The  heat  duty’s  calculated  from  energy  balance  (Stream 12) Heat duty = 1129962.79 kJ/hr + 5% heat loss = 1186460.93 kJ/hr 1186460.93 kJ/hr × 1hr/3600s = 329.572 kW

180 Step 2: Physical Properties Component

C4H10 C4H8-2 C4H6

Mole fraction,x

Surface tension, (dyne/cm) at 34 10.27 11.59 12.78

0.675 0.323 0.002

Average surface tension, Average liquid density, Average vapor density, Average latent heat,

ave L ave

= 11.55 dyne/cm

= 536.22 kg/m

v ave =

ave =

Liquid Vapor Latent density, density, heat (kg/m³) (kg/m³) (kJ/kg) L v at 34 at 110 510.13 1.241 408.726 533.25 1.133 417.5014 565.29 1.069 523.71

1.148 kg/m

449.98 kg/m

Step 3: Overall Coefficient U0 is 1000 W/ .m²  which  is  estimated  from  Figure  12.1  ‘Coulson & Richardson’s,  ‘Chemical  Engineering  Design’  Volume  6  .

Step 4: Exchanger Type and Dimension Tm = True mean temperature difference = 155

- 110

= 45

Step 5: Heat Transfer Area A = Q/U Tm = = 7.324 m

181 Step 6: Layout and Tube Side  This is a kettle reboiler design using U-tube.  Neither fluid is corrosive nor operating pressure is too high, plain carbon steel can be used for the shell and tube.  Saturated dowtherm A will flow inside the tube while the bottom products from distillation column will flow in the shell.  Outside diameter of the tube is chosen as 30mm while inside diameter of the tube is chosen as 25mm. Use square pitch arrangement = 1.5 × (Do of tube) = 1.5(30mm) = 45mm  The nominal length of the tube is 10.2m on a square 45mm pitch.

Step 7: Number of Tubes Area of one tube (neglecting the thickness of tube sheet) A = DL = (30×10-3)(4.8) = 0.4524 m

Number of tubes = = = 16.2

16 tubes

Take minimum bend radius = 1.5

do

= 1.5(30) = 45 mm Total number of tube, Nt= 2×16 = 32 tubes

182 Step 8: Heat Flux, Pool Boiling Coefficent and Overall Heat Transfer Coefficent

Heat flux, q= = = 44.999 kW/m  Checking for maximum allowable heat flux, For square arrangement, Kb = 0.44

qcb = Kb =

0.44

= 156.963kW/m Applying a factor of 0.7, maximum flux should not exceed 0.70×156.963 = 109.87kW/m Actual flux of 44.999 kW/m is well below maximum allowable. Operating pressure: 4.5 bar, Critical pressure = 48.2 bar By  using  Mostinski’s  Equation,  pool  boiling  coefficient,

hnb = 0.104(Pc)0.69(q)0.7 = 0.104(48.2)0.69(0.044999)0.7 = 3912.823 W/m .

183 Taking downtherm A condensing coefficient as 6000 W/m . Downtherm A fouling as 5000 W/m . Bottom product fouling coefficient as 5000 W/m . Conductivity of plain carbon steel ,kw as 45 W/m .

 Calculating Overall Heat Transfer Coefficient

=

+

=

+

+

+

+

×

+

+

×

×

+

×

U0 = 1032.01 Wm .

Error = =

× 100% × 100%

= 3.20%

Therefore, the overall heat transfer coefficient calculated is close enough to the original estimated one for the design to stand.

Step 9: Bundle Diameter and Shell Diameter From Table 12.4, Coulson  &  Richardson’s,  ‘Chemical  Engineering  Design’   Volume 6, for two-tubes passes, K1 = 0.156, n1 = 2.291

184

Db = d0

= 30 = 306.407 mm Taking shell diameter as twice bundle diameter, Ds = 2

306.407

= 612.814 mm

Step 10: Layout and Vapour Velocity Taking liquid level as 400 mm from base, freeboard = 612.814 – 400 = 212.814 mm (satisfactory)

Surface area of liquid = = 0.6837 m

4.6.2 Reboiler (E-105)

 This is a kettle reboiler using U-tube  Do = 30mm, Di = 25mm  Nominal tube length = 10.4 m  Inlet temperature of bottom product = 110  Desired temperature for vapor = 200  Heating medium = Downtherm A, 240

185 Step 1: Specification The  heat  duty’s  calculated  from  energy  balance  (Stream  16a) Heat duty = 1252933.62 kJ/hr + 5% heat loss = 1315580.35 kJ/hr 1186460.93 kJ/hr × 1hr/3600s = 365.43892 kW

Step 2: Physical Properties Component

Mole

Surface

Liquid

Vapor

Latent

fraction,x

tension,

density,

density,

heat

(dyne/cm)

L (kg/m³)

v (kg/m³)

at 40

at 40

at 200

(kJ/kg)

C4H10

0.575

15.57

579

10.64

408.726

C4H8-2

0.425

13.98

604

10.27

417.5014

Average surface tension, Average liquid density, Average vapor density, Average latent heat,

ave L ave

= 14.775 dyne/cm

= 591.5 kg/m

v ave =

ave =

10.455 kg/m

413.11 kg/m

Step 3: Overall Coefficient U0 is 1000 W/ .m²  which  is  estimated  from  Figure  12.1  ‘Coulson & Richardson’s,  ‘Chemical  Engineering  Design’  Volume  6  .

Step 4: Exchanger Type and Dimension Tm = True mean temperature difference = 240 = 55

- 200

186 Step 5: Heat Transfer Area A = Q/U Tm = = 9.13 m

Step 6: Layout and Tube Side  This is a kettle reboiler design using U-tube.  Neither fluid is corrosive nor operating pressure is too high, plain carbon steel can be used for the shell and tube.  Saturated dowtherm A will flow inside the tube while the bottom products from distillation column will flow in the shell.  Outside diameter of the tube is chosen as 30mm while inside diameter of the tube is chosen as 25mm. Use square pitch arrangement = 1.5 × (Do of tube) = 1.5(30mm) = 45mm  The nominal length of the tube is 10.2m on a square 45mm pitch.

Step 7: Number of Tubes Area of one tube (neglecting the thickness of tube sheet) A = DL = (30×10-3)(4.8) = 0.4524 m

Number of tubes = = = 20.2

20 tubes

187 Take minimum bend radius = 1.5

do

= 1.5(30) = 45 mm Total number of tube, Nt= 2×20 = 40 tubes Step 8: Heat Flux, Pool Boiling Coefficent and Overall Heat Transfer Coefficent Heat flux, q= = = 40.026 kW/m  Checking for maximum allowable heat flux, For square arrangement, Kb = 0.44

qcb = Kb = 0.44 = 432.6kW/m Applying a factor of 0.7, maximum flux should not exceed 0.70×432.6.63 = 302.82kW/m Actual flux of 40.026 kW/m is well below maximum allowable. Operating pressure: 4.5 bar, Critical pressure = 40.2 bar By  using  Mostinski’s  Equation,  pool  boiling coefficient, hnb = 0.104(Pc)0.69(q)0.7 = 0.104(40.2)0.69(0.040026)0.7 = 4758.823 W/m .

188 Taking downtherm A condensing coefficient as 6000 W/m . Downtherm A fouling as 5000 W/m . Bottom product fouling coefficient as 5000 W/m . Conductivity of plain carbon steel ,kw as 45 W/m .  Calculating Overall Heat Transfer Coefficient =

+

=

+ +

+ +

×

+ +

× ×

+

×

U0 = 1032.01 Wm .

Error = =

× 100% × 100%

= 3.20% Therefore, the overall heat transfer coefficient calculated is close enough to the original estimated one for the design to stand.

Step 9: Bundle Diameter and Shell Diameter From Table 12.4, Coulson  &  Richardson’s,  ‘Chemical  Engineering  Design’   Volume 6, for two-tubes passes, K1 = 0.156, n1 = 2.291

Db = d0 = 30 = 337.75 mm Taking shell diameter as twice bundle diameter, Ds = 2

337.75

= 675.5 mm

189 Step 10: Layout and Vapour Velocity Taking liquid level as 400 mm from base, freeboard = 675.5 – 400 = 275.85 mm (satisfactory)

Surface area of liquid = = 0.3376 m

4.7 DESIGN OF PUMP

4.7.1 P-102A/B

Calculation Power of Pump Average density = 600 kg/m We take,

= 0.8

(Ref,  Mc  Cabe  Smith  Harriot,”Unit  Operation  of  Chemical  Engineering”  Sixth   Edition) Mass flow rate = 21754.4987 kg/hr

Volume flow rate,v = = = 36.2575 m /hr

190

P (shaft Power)

=

= = 2.1303 kW Power delivered to liquid = 0.8

2.1304

= 1.7042 kW By assuming no velocity difference, Head =

=

5m = 22.21 m

4.7.2 P-103A/B

Calculation Power of Pump Average density = 600 kg/m We take,

= 0.8

(Ref,  Mc  Cabe  Smith  Harriot,”Unit  Operation  of  Chemical  Engineering”  Sixth   Edition) Mass flow rate = 2419.26 kg/hr

Volume flow rate,v = =

191 = 9.0353 m /hr

P (shaft Power)

=

= = 551.459 kW Power delivered to liquid = 0.8 2.1304 = 441.1675 kW By assuming no velocity difference, Head =

=

7m = 45. m

4.7.3 P-104A/B

Calculation Power of Pump Average density = 3615.22 kg/m We take, = 0.8 (Ref, Mc Cabe Smith  Harriot,”Unit  Operation  of  Chemical  Engineering”  Sixth   Edition) Mass flow rate = 2073.86 kg/hr Volume flow rate,v = = = 10.3563 m /hr P (shaft Power)

=

192

= = 608.4784 kW Power delivered to liquid = 0.8 608.4784 kW = 486.7828 kW By assuming no velocity difference, Head =

=

7m = 11.98 m

4.8 DESIGN OF COMPRESSOR

4.8.1 Compressor (C-101A/B)

T=20°C Pinlet= 0.95 atm

T=25°C Poutlet = 1 atm

Component, x

Mole fraction

ρ  @  25oC and 1 atm, (kg/m3),

C4H10

1.0

4.2

Mass flow rate, (kg/hr), mx 6885.3731

Design criteria: The function of this compressor (C-101A/B) is to move n- butane gases stream to the reactor.

193 Design basis: The type of compressor that we used is centrifugal compressor i) Calculate volumetric flow rate, V (m3/s): = 6885.3731 kg/hr

4.2 kg/m3

0.4554 m3/s ii) Calculate the poly tropic efficiency of centrifugal compressor: The equation is referring to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): , Where, = poly tropic efficiency volumetric flow rate, (m3/s) = 0.017 ln (

) + 0.7

= 0.6866 iii) Calculate poly tropic coefficient: n=  γ

/  (γ

γ  +  1),

Where, γ  =  Cp/Cv  =  ratio  of  heat  capacity Refer to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): From  table  13.8,  typical  value  of  γ  for polyatomic gas is 1.30 n = (1.30 x 0.6866)/ (1.30 x 0.6866) – 1.30 +1 n= 1.5063 iv) Calculate the power of centrifugal compressor, W (kW):

W=

194

W=

W= 0.0320 atm.m3/s W=3242.4 J = 3.2424 kW

4.8.2 Compressor (C-102A/B)

T=20°C Pinlet= 0.95 atm

T=25°C Poutlet = 1 atm

Component, x

Mole fraction

ρx @ 25oC and 1 atm, (kg/m3),

O2 N2 H2

0.2853 0.5718 0.1429

1.308 1.251 0.0899

Mass flow rate, (kg/hr), mx 4303.5621 8625.2254 2155.5521

Design criteria: The function of this compressor (C-102 A/B) is to move air stream to the reactor.

195 Design basis: The type of compressor that we used is centrifugal compressor i) Calculate volumetric flow rate, V (m3/s): ρmix = 1.1013 kg/m3 1.1013 kg/m3

V= 15084.3396 kg/hr V= 3.8047 m3/s

ii) Calculate the poly tropic efficiency of centrifugal compressor: The equation is referring to Chemical Process Design and Integration Handbook by

Robin Smith, 2005 (ms274): ,

Where, = poly tropic efficiency volumetric flow rate, (m3/s) = 0.017 ln (3.8047 m3/s) + 0.7 = 0.7229 iii) Calculate poly tropic coefficient: n=  γ

/  (γ

γ  +  1),

Where, γ  =  Cp/Cv  =  ratio  of  heat  capacity Refer to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): From  table  13.8,  typical  value  of  γ  for  diatomic  gas  is  1.40 n = (1.40 x 0.7229/ (1.40 x 0.7229) – 1.40 +1 n= 1.6535 iv) Calculate the power of centrifugal compressor, W (kW):

W=

196 W=

W= 0.2539 atm.m3/s W=25724.6116 J = 25.7246 kW

4.8.3 Compressor (C-103A/B)

T=97°C Pinlet= 0.95 atm

T=100°C Poutlet = 1 atm

Component, x

Mole fraction

ρ @ 100oC and 1 atm, (kg/m3),

H20

1.0

0.596

Mass flow rate, (kg/hr), mx 5341.7169

Design criteria: The function of this compressor (C-103 A/B) is to move steam stream to the reactor.

197 Design basis: The type of compressor that we used is centrifugal compressor

i) Calculate volumetric flow rate, V (m3/s): ρ = 0.596 kg/m3 V= 5341.7169 kg/hr

0.596 kg/m3

V= 2.4896 m3/s ii) Calculate the poly tropic efficiency of centrifugal compressor: The equation is referring to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): , Where, = poly tropic efficiency volumetric flow rate, (m3/s) = 0.017 ln (2.4896 m3/s) + 0.7 = 0.7155 iii) Calculate poly tropic coefficient: n=  γ

/  (γ

γ  +  1),

Where, γ  =  Cp/Cv  =  ratio  of  heat  capacity Refer to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): From  table  13.8,  typical  value  of  γ  for  polyatomic  gas  is  1.30 n = (1.30 x 0.7155/ (1.30 x 0.7155) – 1.30 +1 n= 1.4761 ii) Calculate the power of centrifugal compressor, W (kW):

W=

198 W=

W= 0.1681 atm.m3/s W = 17032.7325 J = 17.0327 kW

4.8.4 Compressor (C-104A/B)

T= 550°C Pinlet= 1 atm

T=560°C Poutlet = 1.10 atm

Mole fraction

ρx @ 550oC and 1 atm, (kg/m3),

C4H10

0.1860

4.2

Mass flow rate, (kg/hr), mx 6885.3731

C4H8-1

0.0614

2.38

2272.9135

C4H8-2

0.1316

2.417

4871.5857

C4H6

0.0127

2.36

470.1303

H2O

0.1443

0.314

531.7169

CO2

0.0065

1.977

240.6179

O2

-

1.308

-

N2

0.2370

1.251

8773.2980

H2

0.2245

0.0899

8310.5713

Component, x

199 Design criteria: The function of this compressor (C-104A/B) is to move mixture of gases stream to the reactor.

Design basis: The type of compressor that we used is centrifugal compressor i) Calculate volumetric flow rate, V (m3/s): ρ mix = 1.4229 kg/m3 1.4229 kg/m3

= 37018.135 kg/hr 7.2265 m3/s

ii) Calculate the poly tropic efficiency of centrifugal compressor: The equation is referring to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): , Where, = poly tropic efficiency volumetric flow rate, (m3/s) = 0.017 ln (

) + 0.7

= 0.7336 iii) Calculate poly tropic coefficient: n=  γ

/  (γ

γ  +  1),

Where, γ  =  Cp/Cv  =  ratio  of  heat  capacity Refer to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): From  table  13.8,  typical  value  of  γ  for  polyatomic  gas  is  1.30 n = (1.30 x 0.7336)/ (1.30 x 0.7336) – 1.30 +1 n= 1.4589

200 iv) Calculate the power of centrifugal compressor, W (kW):

W=

W=

W= 0.9238 atm.m3/s W=93607.3413 J = 93.6073 kW

4.8.5 Compressor (C-105 A/B)

T= 20°C Pinlet= 0.95 atm

T=25°C Poutlet = 1.0 atm

Component, x

Mole fraction

ρx @ 25oC and 1 atm, (kg/m3),

O2 N2

0.21 0.79

1.308 1.251

Mass flow rate, (kg/hr), mx 2605.6283 9802.1257

Design criteria: The function of this compressor (C-105A/B) is to move oxygen gases stream to the reactor.

201 Design basis: The type of compressor that we used is centrifugal compressor i) Calculate volumetric flow rate, V (m3/s): ρ = 1.1854 kg/m3

1.1854 kg/m3

= 12407.754 kg/hr 2.9075 m3/s

ii) Calculate the poly tropic efficiency of centrifugal compressor: The equation is referring to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): , Where, = poly tropic efficiency volumetric flow rate, (m3/s) = 0.017 ln (

) + 0.7

= 0.7181 iii) Calculate poly tropic coefficient: n=  γ

/  (γ

γ  +  1),

Where, γ  =  Cp/Cv  =  ratio  of  heat  capacity Refer to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): From  table  13.8,  typical  value  of  γ  for  diatomic  gas  is  1.40 n = (1.40 x 0.7181)/ (1.40 x 0.7181) – 1.40 +1 n= 1.6608

202 iv) Calculate the power of centrifugal compressor, W (kW):

W=

W=

W= 0.2056 atm.m3/s W= 20828.0861 J = 20.8281 kW

4.8.6 Compressor (C-106A/B)

T=450°C P= 1 atm

T=470°C P= 2 atm

Mass Flow rate(kg/hr) Component C4H10 C4H8-1 C4H8-2 C4H6 H2O CO2 H2 O2 N2

Outlet 7 10052.60487 0 4574.5534 6267.423474 1430.380107 1293.42882 12106.87418 3804.202412 42702.17207

203 Design criteria The function of this compressor (P-106) is to compress the nitrogen, hydrogen, carbon dioxide and oxygen. Design basis  We  know  that  a  pressure  drop  (ΔP)  of  1  atm  in  the  pipeline  from  the   compressor to the discharge point.  The pressure of the fluid at the inlet, P1 = 1atm and pressure of the fluid at the compressor outlet, P2 = 2 atm  The inlet temperature is 450 °C, the calculated outlet temperature is 470°C.

Design Calculations Tout =470° C  For this stream, we assume for diatomic gases which compression ratio = 4, because the specific heat ratio is 1.4 (heuristic rules). Thus, number of stage, n (2/1)1/n = 4 n=  0.5  ≈  1  The mixture of the stream is assume to be diatomic, thus K = 1.4 A= (1.4-1)/1.4 = 0.2857 Power, P P = (3804.202412+42702.17207) kmol /hr ×8.314

× 373K×1.0

× =1170273.04 kJ/ hr =1170.27 kW  From the heuristic, assuming efficiency of 80%. Thus, actual power of the compressor, P actual P actual = =1170.27 / 0.80 = 1462.84kW

204 4.8.6 Compressor (C-107A/B)

T=500°C Pinlet= 1 atm

T=510°C Poutlet = 1.5atm

Component, x

Mole fraction

ρx (kg/m3),

C4H10 N2 H2

0.1057 0.4490 0.1273

2.48 1.251 0.0899

Mass flow rate, (kg/hr), mx 6885.37 29248.18 8292.41

Design criteria: The function of this compressor (C-106) is to move gases stream to the heat exchanger.

Design basis: The type of compressor that we used is centrifugal compressor i) Calculate volumetric flow rate, V (m3/s): = 44425.96 kg/hr

2.48m3/kg

4.976 m3/s ii) Calculate the poly tropic efficiency of centrifugal compressor: The equation is referring to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): ,

205 Where, = poly tropic efficiency volumetric flow rate, (m3/s) = 0.017 ln (

) + 0.7

= 0.7273 iii) Calculate poly tropic coefficient: n=  γ

/  (γ

γ  +  1),

Where, γ  =  Cp/Cv  =  ratio  of  heat  capacity Refer to Chemical Process Design and Integration Handbook by Robin Smith, 2005 (ms274): From  table  13.8,  typical  value  of  γ  for polyatomic gas is 1.30 n = (1.30 x 0.7273/ (1.30 x 0.7273) – 1.30 +1 n= 1.4648 iv) Calculate the power of centrifugal compressor, W (kW):

W=

W=

W= 0.260304 atm.m3/s W=26030 W= 26.03 kW

206 4.9 DESIGN OF STORAGE TANK

4.9.1 Storage Tank TK-1

Design criteria To collect butadiene from the separation process before it goes to consumer. The tank is design for the volume of 2 hours for the inlet mass flow. Design Parameter  Volume of the storage tank  Diameter and height of the tank  Type of the tank

Design Basis  Total inlet flow rate : 6250 kg/hr  Temperature

: 30

 Pressure

: 1.0 atm

Design Calculation ρbutadiene = 54 kg/m Volumetric flow rate = = = 115.74 m /hr

207 Retention time = 2 hours Volume of the product for 2 hours = 115.74 m /hr × 2h = 231.48 m The storage tank is designed as 25% bigger than the product volume. V= (1.25) × (231.48) = 289.35 m Assume L/D = 3

=

× =

×3

D = 122.8 D = 4.97 m L=D×3 = 4.97 × 3 = 14.91 m So, the diameter of storage tank is 4.97m and height is 14.91m. The storage tank is designed as horizontal tank on concrete supports.

4.9.2 Storage Tank TK-2 Design criteria To collect 1-butene from the separation process before it goes to consumer. The tank is design for the volume of 2 hours for the inlet mass flow. Design Parameter  Volume of the storage tank  Diameter and height of the tank  Type of the tank

208 Design Basis  Total inlet flow rate : 7117.66732 kg/hr  Temperature

: 34

 Pressure

: 1.0 atm

Design Calculation Ρ1-Butene = 600 kg/m Volumetric flow rate = = = 11.863 m /hr Retention time = 2 hours Volume of the product for 2 hours = 11.863 m /hr × 2h = 23.726 m The storage tank is designed as 25% bigger than the product volume. V= (1.25) × (23.726) = 29.656 m Assume L/D = 3

=

× =

D = 12.58 D = 2.33 m L=D×3 = 2.33 × 3 = 6.978 m

×3

209 So, the diameter of storage tank is 2.33m and height is 6.978m. The storage tank is designed as horizontal tank on concrete supports.

210

CHAPTER 5

PROCESS SAFETY STUDIES

5.1

Introduction Safety has to be an integral part of any process plant. Safe facilities that

meet performance and economic requirements are in the best interest of all concerned parties. Unless personnel know and understand the principles behind process and plant safety, there is an increased risk of accidents that could result in injuries, illness, fatalities and loss property. The importance of a safety review is being recognized as an important risk management tool. Failure to identify risks to safety, and the according inability to address or "control" these risks, can result in massive costs, both human and economic. The multidisciplinary nature of safety engineering means that a very broad array of professionals is actively involved in accident prevention or safety engineering. The majority of those practicing safety engineering are employed in industry to keep workers safe on a day to day basis All manufacturing processes are to some extent hazardous but in chemical processes, there are additional, special hazards associated with the chemicals used and the process condition. It is very important to ensure the risks and hazards of chemical processes are reduced to acceptable level. Therefore identifying all potential hazards can control accidents and hazards.

In this chapter, general process safety will be discussed. Hazard and operability (HAZOP) studies also discussed to identify the potential hazards associated with operation of the plant and consequently, the appropriate actions to be taken. Appraisals of plant safety should be prepared as long before construction

211 is started. This mean, the safety aspects should be considered as soon as new plant is conceived or design begun. The plant safety is managed by taken many parties into consideration. Generally, it is conducted by plant safety personnel, safety committee members, and plant or department managers and together with other related department or agencies include fire prevention personnel, insurance company engineers or municipal, state, or federal agency representatives. Checklists are often used to evaluate many safety features in industrial plants which represented as question asking the specific requirements in the standards. It can be prepared from numerous sources concern on to determine whether or not a plant can comply with Occupational Safety and Health Act (OSHA) standards.

The chemical industry has traditionally dedicated considerable attention to safety, beginning with the research and development new processes through plant design and construction. In chemical plant, the main hazards are toxic and corrosive chemicals, explosions, fires and accidents common to all industrial activities. The Health and Safety at Work (Act 1947) provided a new legal administrative framework to promote, simulate and encourage even higher standards of health and safety at work.

It has been found that after many years of improvements in technical safety methods and process design, many organizations have found that accident rates, process plant losses and profitability have reach a level beyond which further improvements seem impossible to

achieve. Another findings are that

even in organizations with good general safety records, occasional large scale disasters occurs which shake public confidence in the chemical process industry (Guidelines for preventing human error in Process Safety, 1994).

Besides, some hazardous in chemical plant processes are still to be avoided by all the manufacturing process (R.K Sinnott et el, 1983). In this chapter of process safety studies as well as special hazards of chemicals shall be reviewed. Also in this chapter the procedure to start-up and shutdown in general will be considered.

212 5.2

Identification of Hazard

The hazard identification methods include process hazard checklist, hazards surveys, HAZOP and safety review. One of the major characteristics of the HAZOP study is to identify potential hazards such as the corrosive and toxicity. This is essential to identify the hazards and reduce the risk well in advance of an accident as well as a guideline to protect the worker.

5.2.1 Corrosive

A Corrosive material is one which causes damage to skin, eyes or other parts on the body on contact. The technical definition is written in terms of "Destruction or irreversible damage to living tissue at the site of contact"

Often this damage is caused directly by the chemical, but the action of some corrosive materials is a consequence of inflammation which they may cause. Concentrated acids are obvious examples of corrosive materials, but even dilute solutions of bases such as sodium or ammonium hydroxide may also be very corrosive, particularly in contact with the eyes.

5.2.2 Toxicity

The potential hazard will depend on the inherent toxicity of the material and the frequency and duration of any exposure. The inherent toxicity of a material is measured by tests on animals which are usually expressed as the lethal dose (LD50) at which 50% of the test animals are killed. These definitions apply only to the short-term   effects.   Another   symbol   describing   toxicity   is   “Threshold   Limit   Value”   (TLV)   for controlling long-term exposure of workers to contaminated air. Toxicity is often subdivided into:

213  Acute toxicity adverse effects are observed within a short time of exposure to the chemical. This exposure may be a single dose, or a short continuous exposure, or multiple doses administered over 24 hours or less.  Subacute (subchronic) toxicity adverse effects are observed following repeated daily exposure to a chemical, or exposure for a significant part of an organism's lifespan (usually not exceeding 10%). With experimental animals, the period of exposure may range from a few days to 6 months.  Chronic toxicity adverse effects are observed following repeated exposure to a chemical during a substantial fraction of an organism's lifespan (usually more than 50%). For humans, chronic exposure typically means several decades; for experimental animals, it is typically more than 3 months. Chronic exposure to chemicals over periods of 2 years using rats or mice may be used to assess the carcinogenic potential of chemicals.

5.2.3 Flammable

The flammability term is to describe materials that will burn. Flammable material depends on a number of factors:

5.2.3.1

Flash point

The flash-point is a measure of the ease of ignition of the liquid. It is a function of the vapor pressure and the flammability limits of the material.

214 5.2.3.2

Flammability limits

The flammability limits of a material are the lowest and highest concentrations in air, at normal pressure and temperature, at which a flame will propagate through the mixture. It is differ widely for different materials.

5.3

Control of Hazard Analysis

5.3.1 1,3-Butadiene

1,3-Butadiene is a product that we get from oxidative dehydrogenation of n-butene which is colorless gas with a mild, aromatic, gasoline-like odor. The vapor is heavier than air and commercially available as a liquefied gas (under pressure) with a stabilizer added for shipment. It is an important industrial uses as monomer in the production of synthetic rubber. It is colorless gas or refrigerated liquid with boiling point -4.4ºC. It needs to be storage in area below 40ºC. 1,3butadiene always be marketed and transported as a liquefied gas under pressure.

Health Concerns: 

At acute exposure, damage to the central nervous system will occur. Symptoms such as distorted blurred vision, vertigo, general tiredness, decreased blood pressure, headache, nausea, decreased pulse rate and fainting may be witnessed. Several studies show that 1,3-Butadiene exposure increases risk in cardiovascular diseases and cancer. It also can cause irritation with redness to skin and eyes. Exposure to liquid may cause frostbite.



While handling this substance, we need to take precaution step as protections. Whenhandling the substance, the person in-charged needs to use chemical protective clothing(CPC) like polyvinyl chloride gloves to prevent contact

215 with chemicals that can injure or be absorbed through the skin. The level of protection selected should be based on the potential butadiene concentration and likelihood of contact.

Safety concerns: 

1,3-Butadiene is non-corrosive but highly flammable substance and can forms explosive mixtures with air and oxidizing agents. It also sensitive to impact, so container should be avoid from impact. Its extinguish media are CO2, dry chemical, water spray or fog. As precautions, the substances need to be store and use with adequate ventilation. The storage vessel should be enclosed within fire banks.



All vessels for storage and handling liquefied the substances should meet recognized code requirements, such as ASME Code for Unfired Pressured Vessels. The vessels and pipes should be the quality of solid masonry, concrete or steel. If steel is used, it should be protected against fire exposure. All parts of the storage system should be electrically connected to one another and to a common ground in a way that will prevent the accumulation of static electrical charge.



Pressure relief valves should be connected to a burning stack or flare system provided with a positive means of ignition. Totally enclosed gauging devices generally should be used.

Environmental concerns: 

1,3-butadiene is an unstable chemical substance and spontaneously flammable or explosive peroxides when exposure to air. Thermal decomposition or burning may produce carbon monoxide or carbon dioxide.

216 

No adverse ecological effect is expected. This product does not contain any Class I or Class II ozone-depleting chemicals. The components of this mixture are not listed as marine pollutants by TDG Regulations too.

5.3.2 n-Butane

n-butane which is consists of 1-butene, cis-2-butene and trans-2butene are being used as raw material in the oxidative dehydrogenation process to produce 1,3-butadiene. However, only 1-butene is being discussed. Its boiling point is 20.8ºC (-6.25ºF) with vapor pressure 263.3 kPa at 20ºC.

Health Concerns: 

At overexposure, it can cause asphyxiant and death due to the lack of oxygen. Moderateconcentration of 1-butene may cause headaches, drowsiness,

dizziness,

excitation,

excess

salivation,

vomiting

and

unconsciousness. It also can cause frostbite to skin and eyes when it is in liquid phase. 

Precaution should be taken when handling the substance. Handler should use respirable fume respirator or air supplied respirator when working in confined space or where local exhaust or ventilation does not keep exposure below TLV. They also need to wear safety gloves and glasses when handling cylinders that contain 1-butene.

Safety concerns: 

The substances are flammable and high pressure gas. It may form explosive when mixing with air. The cylinders should be protected from damage by using suitable hand truck to move the cylinders; do not drag, roll, slide or

217 drop it. Never attempt to lift a cylinder by its cap; the cap is intended solely to protect the valve. 

Never insert an object into cap openings; doing so may damage the valve and cause a leak. Precautions should be taken also in storage the substance. The cylinders should be separated from oxygen, chlorine, and other oxidizers material. Firmly secure cylinders upright to keep them from falling or being knocked over. There must be no sources of ignition and all electrical equipment in storage areas must be explosion –proof. Storage areas must in temperature below 52ºC.

Environmental concerns: 

1-butene does not contain any Class I or Class II ozone-depleting chemicals.



The components of this mixture are not listed as marine pollutants by TDG Regulations.

5.3.3 n-methyl-2-pyrrolidone

n-Methyl-2-Pyrrolidone (NMP) is a solvent that is being used in purification process ofcrude butadiene. It is colorless hygroscopic liquid with characteristic odor. Its boiling point is 202ºC and flash point is 96ºC.

Health Concerns: 

NMP also can cause dry, redness and maybe absorbed to skin. As precaution, useprotective gloves and clothing. When handling the substance, safety spectacles should be used to prevent eyes from redness, pain and blurred vision. Avoid from eating, drinking or smoking during work.

218 Safety concerns: 

Because of it is flammable substance, storage area should be separated from oxidants, rubber, plastics, aluminium and light metal.



Make sure it place in dry and ventilation along the floor. Above 96ºC, explosive vapor/air mixture may be formed.



Therefore, open flames are prohibits and use a closed system with ventilation when handle it in temperature above 96ºC.

Environmental concerns: 

It will turn yellow on exposure to heat. This substance decomposes on heating or onburning producing toxic fumes including nitrogen oxides, carbon monoxide. NMP is combustible substance and can give off irritating or toxic fumes in fire.

5.3.4 Dow Fire and Explosion Index The hazard classification guide developed by the Dow Chemical Company and published by the American Institute of Chemical Engineering, Dow (1994) (www.aiche.org), gives a method of evaluating the potential risk from a process, and  assessing  the  potential  loss.  A  numerical  “Fire  and  explosion  index”  (F  &  EI)   is calculated, based on the nature of the process and the properties of the process materials. The larger the value of the F & EI, the more hazardous the process, see Table 5.1.

219

Table 5.1: Assessment of hazard (Adapted from the Dow F & EI guide 1994) To assess the potential hazard of a new plant, the index can be calculated after the Piping and Instrumentation and equipment layout diagrams have been prepared. In earlier versions of the guide the index was then used to determine what preventative and protection measures were needed, see Dow (1973). In the current version the preventative and protection measures that have been incorporated in the plant design to reduce the hazard are taken into account when assessing the potential loss; in the form of loss control credit factors. It is worthwhile estimating the F & EI index at an early stage in the process design, as it will indicate whether alternative, less hazardous, process routes should be considered.

Only a brief outline of the method used to calculate the Dow F & EI will be given in this section. The full guide should be studied before applying the technique to a particular process. Judgment, based on experience with similar processes, is needed to decide the magnitude of the various factors used in the calculation of the index, and the loss control credit factors.

5.3.4.1

Calculation of the Dow F & EI The first step is to identify the units that would have the greatest impact on

the magnitude of any fire or explosion. The index is calculated for each of these units. The basis of the F & EI is a Material Factor (MF). The MF is then multiplied by a Unit Hazard Factor, F3, to determine the F & EI for the process

220 unit. The Unit Hazard factor is the product of two factors which take account of the hazards inherent in the operation of the particular process unit: the general and special process hazards.

5.3.4.2 Material factor

The material factor is a measure of the intrinsic rate of energy release from the burning, explosion, or other chemical reaction of the material. Values for the MF for over 300 of the most commonly used substances are given in the Figure 5.1. The Figure5.1 also includes a procedure for calculating the MF for substances not listed: from knowledge of the flash points, (for dusts, dust explosion tests) and a reactivity value, Nr. The reactivity value is a qualitative description of the reactivity of the substance, and ranges from 0 for stable substances, to 4 for substances that are capable of unconfined detonation. Some typical material factors are given in Table 5.2. In calculating the F & EI for a unit the value for the material with the highest MF, which is present in significant quantities is used.

221

Figure 5.1: Procedure for calculating the fire and explosion index and other risk analysis information.

222

Table 5.2: Some typical material factors

5.3.4.3 General process hazards The general process hazards are factors that play a primary role in determining the magnitude of the loss following an incident. Six factors are listed on the calculation form, Figure5.2. A.

Exothermic chemical reactions: the penalty varies from 0.3 for a mild exothermic, such as hydrogenation, to 1.25 for a particularly sensitive exothermic, such as nitration.

B.

Endothermic processes: a penalty of 0.2 is applied to reactors, only. It is increased to 0.4 if the reactor is heated by the combustion of a fuel.

C.

Materials handling and transfer: this penalty takes account of the hazard involved in the handling, transfer and warehousing of the material.

D.

Enclosed or indoor process units: accounts for the additional hazard where ventilation is restricted.

223 E.

Access of emergency equipment: areas not having adequate access are penalized. Minimum requirement is access from two sides.

F.

Drainage and spill control: penalizes design conditions that would cause large spills of flammable material adjacent to process equipment; such as inadequate design of drainage.

224

Figure 5.2:

Dow Fire and Explosion Index Calculation Form

225 5.3.4.4 Special Process Hazards

The special process hazards are factors that are known from experience to contribute to the probability of an incident involving loss. Below are twelve factors taking consideration on the calculation:

A.

Toxic materials: the presence of toxic substances after an incident will

make the task of the emergency personnel more difficult. The factoapplied ranges from 0 for non-toxic materials, to 0.8 for substances that can causedeath after short exposure. B.

Sub-atmospheric pressure: allows for the hazard of air leakage into

equipment. It is only applied for pressure less than 500 mmHg. C.

Operation in or near flammable range: covers for the possibility of air

mixing with material in equipment or storage tanks, under conditions where the mixture will be within the explosive range. D.

Dust explosion: covers for the possibility of a dust explosion. The degree

of risk is largely determined by the particle size. The penalty factor varies from 0.25 for particles above 175m, to 2.0 for particles below 75m. E.

Relief pressure: this penalty accounts for the effect of pressure on the rate

of leakage, should a leak occur. Equipment design and operation becomes more critical as the operating pressure is increased. The factor to apply depends on the relief device setting and the physical nature of the process material. F.

Low temperature: this factor allows for the possibility of brittle fracture

occurring in carbon steel, or other metals, at low temperatures. G.

Quantity of flammable material: the potential loss will be greater the

greater the quantity of hazardous material in the process or in storage. The factor

226 to apply depends on the physical state and hazardous nature of the process material, and the quantity of material. It varies from 0.1 to 3.0. H.

Corrosion and erosion: despite good design and materials selection, some

corrosion problems may arise, both internally and externally. The factor to be applied depends on the anticipated corrosion rate. The severest factor is applied if stress corrosion cracking is likely to occur. I.

Leakage joints and packing: this factor accounts for the possibility of

leakage from gaskets, pump and other shaft seals, and packed glands. The factor varies from 0.1 where there is the possibility of minor leaks, to 1.5 for processes that have sight glasses, bellows or other expansion joints. J.

Use of fired heaters: the presence of boilers or furnaces, heated by the

combustion of fuels, increases the probability of ignition should a leak of flammable material occur from a process unit. The risk involved will depend on the sitting of the fired equipment and the flash point of the process material. K.

Hot oil heat exchange system: most special heat exchange fluids are

flammable and are often used above their flash points; so their use in a unit increases the risk of fire or explosion. The factor to apply depends on the quantity and whether the fluids are above or below its flash point L.

Rotating equipment: this factor accounts for the hazard arising from the

use of large pieces of rotating equipment: compressors, centrifuges, and some mixers.

5.3.5 Potential loss The first step is to calculate the Damage factor for the unit. The Damage factor depends on the value of the Material factor and the Process unit hazards factor. An estimate is made of the area (radius) of exposure. This represents the

227 area containing equipment that could be damaged following a fire or explosion in the unit being considered. An estimate of the replacement value of the equipment within the exposed area is then made, and combined with by the damage factor to estimate the Base maximum probable property damage (Base MPPD). The Maximum probable property damage (MPPD) is then calculated by multiplying the Base MPPD by a Credit control factor. The Loss control credit control factors, allow for the reduction in the potential loss given by the preventative and protective measures incorporated in the design. The MPPD is used to predict the maximum number of days which the plant will be down for repair, the Maximum probable  day’s  outage (MPDO). The MPDO is used to estimate the financial loss due to the lost production: the Business interruption (BI). The financial loss due to lost business opportunity can often exceed the loss from property damage. (Chemical Process Safety, Daniel A. Crowl/Joseph F. Louvar, 2002)

5.3.6 Basic Preventative and Protective Measures The basic safety and fire protective measures that should be included in all chemical process designs are listed below. This list is based on that given in the Dow Guide, with some minor amendments.

1. Adequate, and secure, water supplies for fire fighting. 2. Correct structural design of vessels, piping, steel work. 3. Pressure-relief devices. 4. Corrosion-resistant materials, and/or adequate corrosion allowances. 5. Segregation of reactive materials. 6. Earthing of electrical equipment. 7. Safe location of auxiliary electrical equipment, transformers, switches gear. 8. Provision of back-up utility supplies and services. 9. Compliance with national codes and standards. 10. Fail-safe instrumentation. 11. Provision for access of emergency vehicles and the evacuation of personnel.

228 12. Adequate drainage for spills and fire-fighting water. 13. Insulation of hot surfaces. 14. No glass equipment used for flammable or hazardous materials, unless no suitable alternative is available. 15. Adequate separation of hazardous equipment. 16. Protection of pipe racks and cable trays from fire. 17. Provision of block valves on lines to main processing areas. 18. Protection of fired equipment (heaters, furnaces) against accidental explosion and fire. 19. Safe design and location of control rooms.

5.4

Major Equipment Control

This plant consist six major equipments. There are three reactor and three distillation columns. Each of these equipments has their own hazard. Below is the equipment process control for each equipment. The primary objectives of the instrumentation and control schemes are:-

(i)

Safety of plant operation:

To keep the process variables within known safe operating limits To detect dangerous situations as they develop and to provide alarms and automatic shut-down system. (ii)

Production rate: -To achieve the design product output

(iii)

Product Quality: - To maintain the product within the specified quality

standards. (iv)

Cost: - To operate at the lowest production cost, commensurate with the

other objectives.

229 5.4.1

Reactor Control System

FT

FIC

PIC

GAS INLET PT TT

TIC COLD WATER

E-1

P-31

HOT WATER

OUTLET

Figure 5.3: Control System of Reactor (R-101)

The control of chemical reactor is the most important in chemical engineering and must be carefully consent, since the chemical reactor is the main of the processes. The stability efficiency of reaction operation is very important to the entire plant. The reaction employed in the plant is exothermic and runaway operation cause explosion. Thus, it requires an effective, well-designed control system to assure stable operation. The control system objectives are defined as: (i) To maintain the reactor operation at optimum temperature and pressure.

230 (ii) To sustain the product quality in order to make sure the maximum production rate can be maintained at the annual production rate of 50,000 metric ton per year.

(iii)

Provide the safeguards against the unexpected process runaway and maintain the safe operation.

(iv)

To maintain optimum feed condition and composition in order to ensure

high production rate.

Objective

Control Manipulat Disturbanc Variable ed e Variable

Type of Control

Set Point

To maintain the flow rate of inlet and outlet

Inlet flow rate valve

Inlet flow rate

Outlet flow rate

Cascade (master)

37 018.135 kg/hr

To maintain the pressure in the reactor

Inlet flow rate valve

Inlet flow rate

Outlet flow rate

Cascade (slave)

1 atm

To maintain the temperature of the reactor

Input of cold water valve

Input of hot water

Outlet flow rate

Feedback controller

Not below or exceed than 10% of temperature set point which is 550°C

Table 5.3: Control Mechanism of Multi-tube Fixed bed Reactor

231 5.4.2

Distillation Column Control System

PT

D-101

LIC

PIC LT AIC

E-103

AT

TK-1

Stream 11

Reflux Drum

Stream 13 FT TT FIC TIC

NMP AIC

Steam LT

E-104 LIC

AT

Stream 12

P-113A/B

Figure 5.4: Control System of Extractive Distillation Column (D-101)

The control system of distillation column has three main factors to consider: 

Material balance control.



Product quality control.



Satisfaction of constraints.

232 5.4.2.1 Material Balance Control

The column control system must cause the average sum of the product streams (bottom and top product) to be exactly equal to the average feed rate, keeping the column in material balance. Although the plant is usually designed for a nominal production rate, a design tolerance is always incorporated because the market condition and demand may require an increase or decrease from the current state. The control system is then called to ensure a smooth and safe transition from the old production level to the newly desired production level. Its purpose is to direct the control action in such a way as to make the inflows equal to the outflows and achieve a new steady-state material balance for the plant.

5.4.2.2 Product Quality Control

The purpose of this control is to maintain the desired concentration of products at bottom and top stream.

5.4.2.3 Satisfaction of Constrains

For safety purposes, satisfactory operation of the column and certain constraints must be understood and followed, for example, the column shall not flood. Column pressure drop should be high enough to maintain effective column operation, that is, to prevent serious weeping or dumping. The temperature difference in the reboiler should not exceed the critical temperature difference. Avoid shock loading to the column to avoid overload reboiler or condenser heat-transfer capacity.

233 Column pressure should not exceed a maximum permissible value. The main purpose of this distillation column, is to separate acetic acid from the rest of the feed, and it withdraws from the bottom. Objectives:



Control the top pressure in the column



Control the top temperature in the column



Control the temperature in the condenser



Control the liquid level in the condenser



Control the liquid level in the column



Control the product purity Control

Measured

Manipulated

Variables

Variables

Variables

Column

Pressure at the top

Stream 11

Change in

Pressure

tray

flow rate

pressure column

Stream flow

Change in

to the boiler

temperature

Temperature in Temperature at the column

top tray

Disturbances

Type of Controller Feedback Feedback

column Liquid level in

Liquid level in

Flow rate of

Change in liquid

condenser

reflux drum

reflux stream

level in the

Feedback

condenser Flow rate of

Flow rate of

Stream 11

Change in flow

1) Feedback

streams

1) Stream 13

flow rate

rate of streams

2) Feedback Feedback

2) Stream NMP Liquid level at

Liquid level at the

Flow rate of

Change of liquid

the bottom

bottom tray

Stream12

level at the bottom

Qualitity of

Composition in

Flowrate of

Change of

composition in

Stream 13

Stream 13

composition in

the product

the strean of

(Stream 13)

product

Table 5.4: Control Mechanism of Extractive Distillation Column

Cascade

234 5.5

Hazard and Operability Studies (HAZOP)

The Hazard and Operability Study is a procedure for the systematic, critical, examination of the operability of a process. When applied to a process design or an operating plant, it indicates potential hazards that may arise from deviations from the extended design conditions. HAZOP include forms of study, systemic inspection for process plant to determine the danger, probability of failure and problem that occur and consequences if failure occur. HAZOP studies are performed by a team consisting of plant operators, engineers, managers and others, some of whom should be intimately familiar with the facility being studied.

5.5.1 HAZOP Concept

The HAZOP process is based on the principle that a team approach to hazard analysis will identify more problems than when individuals working separately combine results. The HAZOP team is made up of individuals with varying backgrounds and expertise. The expertise is brought together during HAZOP sessions and through a collective brainstorming effort that stimulates creativity and new ideas, a thorough review of the process under consideration is made.

5.5.2 HAZOP objectives

HAZOPs concentrate on identifying both hazards as well as operability problems. The objectives of HAZOP study are:

235 

To identify areas of the design that may possess a significant hazard potential.



To familiarize the study team with the design information available.



To ensure that a systematic study is made of the areas of significant hazard potential.



To determine all deviations that most might happen and make suggestion to minimize it.

5.5.3 Important Equipments for HAZOP 

Process diagram and instrumentation (P&ID)



Mass and energy balances



Safety procedures documents



Involved chemicals



Piping specification



Previous HAZOP report

5.5.4 HAZOP planning 

Time available



Time assumption needed



Available of team members



Available of P&IDs



Report of record



Actions that should be take



Last report and date

236 5.5.5 HAZOP Team

Ideally, the team consists of five to seven members, although a smaller team could be sufficient for a smaller plant. If the team is too large, the group approach fails. On the other hand, if the group is too small, it may lack the breadth of knowledge needed to assure completeness. The team leader should have experience in leading a Hazop. The rest of the team should be experts in areas relevant to the plant operation. A team might include:

a)

Design engineer

b

Process engineer

c)

Operations supervisor

d)

Instrument design engineer

e)

Chemist

f)

Maintenance supervisor

g)

Safety engineer (if not Hazop leader)

The team leader‟s most important job is to keep the team focused on the key task: to identify problems, not necessarily to solve them. There is a strong tendency for engineers to launch into a design or problem-solving mode as soon as a new problem comes to light. Unless obvious solutions are apparent, this mode should be avoided or it will detract from the primary purpose of Hazop, which is hazard identification. In addition, the team leader must keep several factors in mind to assure successful meetings: (1) do not compete with the members; (2) take care to listen to all of the members; (3) during meetings, do not permit anyone to be put on the defensive; (4) to keep the energy level high, take breaks as needed.

237 5.5.6 HAZOP Procedure

Use the following steps to complete an analysis of HAZOP:

a)

Begin with a detailed flow sheet and break the flow sheet into a number of process units. After that select a unit to study

b)

Choose a study node (vessel, line, operating instruction).

c)

Describe the design intent of study node.

d)

Pick a process parameter (level, flow, temperature, pressure, volume, reaction, and component)

e)

Apply a guide word to the process parameter to suggest possible deviation

f)

Determine possible causes and note any protective systems.

g)

Evaluate the consequences of the deviation (if any)

h)

Recommend action

i)

Record all information.

Guidewords No or not

Meanings

Comments

The complete negation of

No part of the intentions

these intentions

is achieved but nothing else happens

More (more of, higher)

Quantitative increases

These refer to the quantities and properties such as flow rates and temperatures as well as activities such as heat and react.

Less (less of, lower)

Quantitative decreases

These refer to the quantities and properties

238 such as flow rates and temperatures as well as activities such as heat and react. As well as

A qualitative increase

All the design and operating intentions are achieved together with some additional activities.

Part of

A qualitative decrease

Only some of the design intentions are achieve; some are not

Reverse

The logical opposite of

This is mostly applicable

the intention

to activities for example reverse flow or chemical reactions.

Other than

Complete substitution

No part of the original intention is achieved. Something quite different happens.

Table 5.5: List of HAZOP Guide Words

239 5.5.7 HAZOP on Storage Tank Deviation

Causes

Consequences

Existing

Actions,

Provision

Questions, or Recommend ations

More flow

Less flow

No Flow

- Control

- Settling of the

- Flow

- Sufficient

valve

liquid phase

indicator and

provision

failure

not

control

effective

- High flow

- Flooding in

alarm,

the tank

HFA

- Pipe line

- Loss of

- Flow

- Sufficient

leakage

production

indicator and

provision

- Control

yield

control

valve

- Low flow

partially

alarm,

closed

LFA

- Pipeline

- Settling tank

- Flow

- Sufficient

rupture

dries

indicator and

provision

- Blocking in

up

control

inlet

- Separation

- Low flow

pipeline

between

alarm,

- Control

products fail

LFA

valve fails to open Table 5.6: HAZOP Studies on Storage Tank

240 5.5.8 HAZOP on Heat Exchanger Deviation

Causes

Consequences

Existing

Actions,

Provision

Questions, or Recommendations

Cold Flow

-Pipe leakage

-Desire

-Install a filter at

-Pump not

temperature

the

function

not

Coolant

accomplished

-Install flow

-Effect further

indicator

process

-Install temperature indicator

Excess flow

-Valve not

Same  as  ‘  Cold

-Install an orifice

of the

Functioning

Flow’

plate

coolant

-The coolant

-Install a cooling

pump

water

cannot be

flow meter

control,

-Install flow

hence the

indicator

capacity

Change new valve

rate increase -The coolant pipe rupture Less flow of

-Control

Same  as  ‘  Cold

-Install a less flow

the coolant

valve

Flow’

warning device

malfunction

-Change a new

-Flow pipe of

control

the

valve

coolant plugged

241 Higher

-Heat

-Pressure

-Install high

Temperature

exchanger is

increase

temperature

malfunction

-Explosion can

warning device at

-Temperature

occur

the out flow of hot

of the hot

-Failure of the

fluid stream

fluid is too

equipment

-Repair the damage

high

of

-Fire beside

the heat exchanger

the heat exchanger Lower

-No

Breakdown of

-Install temperature

temperature

temperature

next

low

Change

equipment

alarm at the out

-

flow

Condensation

-Clean the tube of

happen at the

the

tube of

heat exchanger to

heat

removes deposit

exchanger Table 5.7: HAZOP Studies on Heat Exchanger

242 5.5.9 HAZOP on Multi-tube Fixed Bed Reactor

Deviation

Causes

Consequences

Existing

Actions,

Provision

Questions, or Recommendation s

No

More

No flow

i. Tube leakage

i. No reaction

i. Emergency

and blocking

ii. No

shutdown

ii. Valve fails or

production

ii. Install flow

malfunctioning

iii Disturbing

indicator in inlet

iii. No flow from

others process

flow

previous

iii. Checking

equipment

stream condition

Higher feed i. Valve cannot be flow

. Rise in

i. Install flow

controlled

bottom liquid

indicator

normally

level and

ii. Regular

ii. Sudden

pressure

inspection

increment in fresh

ii. Causes

feed

flood in the reactor iii. Reduce conversion

Higher

i. Less process

i. Fracture in

i. Install

pressure

fluid flow into the

transfer line

temperature alarm

and

heater

due by

high indicator

temperature

ii. More Steam

overload

ii. Proper process

entering heater

temperature

control at heater

ii. Coke deposits form on the catalyst iii. Reduce the

243 conversion Less

Less flow

i. Less flow from

i. Less

i. Install low flow

previous

production

alarm and leakage

equipment

ii.Temperatur

alarm

ii. Leakage or

e drops,

ii. Regular

blockage at piping

causing less

inspection

or no iii. Reactions occurs iv. Causing reversible flow to reactor Low

i. The reactor

i. Yield and

i. Check TIC and

pressure

pressure and

efficiency of

PIC controller

and

ii. temperature

products

ii. Install alarm

decreases

iii. Monitor

ii. Backflow

pressure at feed

from reaction

inlet

temperature iii.malfunctioning

Reverse

Reverse

i. High pressure in

i. No

Install  „check  

flow either

column

operation

valve'

to reactor

ii. Failure in PIC

ii. No

or to feed

iii. Pipe leakage

production

storage

iii. Plant break down

As well

Other flow

i. Contaminants

Quality of

as

of

detected

product

substances

ii. Non-pure feed

decreases

detected for

iii. Leakage in

pureness of feed

example

cooling system

ii. Check the

i. Check the

contaminati on of cool

composition

water

controller

244

Other

Production

i. Leakage of

Affect the

Check and correct

than

of

ii. Heat exchanger

reaction

the reactor

substances

and

other than

its substances

the product. entered the reactor

Table 5.8: HAZOP Studies on Reactor

internals

245 5.5.10 HAZOP on Extractive Distillation Column

HAZOP for Feed Stream Deviation

Possible

Possible

Existing

Action

cause

consequences

provision

required

No Feed

-Blockage of

-Level decrease in

-Install flow

flow

leakage

the column

indicator

-Control

-Loss in

controller (FIC)

system

production

-Check and

damage

recheck/ maintenance

More feed

-Control

-Damage/broken

-Install flow

flow

damage

of column

indicator

failure

-High pressure

controller (FIC)

flow thus increase level in column More

-Pressure

-Stream line

-Pressure

-Place valve at

pressure

valve failure

fracture

indicator

critical

controller

instrument list

-

-Sufficient provision

-Thermal expansion in stream line More

-Temperature

-Line fracture

temperature

built up in

Temperature

stream line

indicator controller (TIC)

Less flow

Same as NO

Same as NO or

Same as NO or

or NOT

NOT

NOT

Less

-Unstable

-Sudden change in

-Same for

pressure

flow

pressure drop

MORE for

condition in

pressure

246 column

condition

Less

-Temperature

-Loss product in

-Same for

temperature

decrease in

upstream

MORE

column

condition

system

HAZOP for Top Upstream

No flow

-Control system

-Loss in

-Install

damage

production

flow indicator controller (FIC)

More flow

-Control

-Damage/

-Install

damage failure

broken of

flow

column

indicator controller (FIC)

Less flow

Same as NO or

Same as NO or

Same as

NOT

NOT

NO or NOT

HAZOP for Bottom Stream No flow

-Control system

-Loss in

-Install

damage

production

flow indicator controller (FIC)

More flow

-Control

- Damage/

-Install

damage failure

broken of

flow

247 column

indicator controller (FIC)

Less

Same as NO or

Same as NO or

Same as

NOT

NOT

NO or NOT

Table 5.9: HAZOP Studies on Distillation Column

248

CHAPTER 6

WASTE MANAGEMENT

6.0

Introduction

Waste management is the collection, transport, processing, recycling or disposal, and monitoring of waste materials (Anonymous, 2009). The term usually relates to materials produced by human activity and is generally undertaken to reduce their effect on health, environment or aesthetics. Waste management is also carried out to recover resources from it. Furthermore, waste management can be involving solid, liquid, gaseous or radioactive substances, with different methods and fields of expertise of each.

Waste management practices differ for developed and developing nations, for urban and rural areas, and also for residential and industrial producers. Management for non-hazardous residential and institutional waste in metropolitan areas is usually the responsibility of local government authorities, while management for non-hazardous commercial and industrial waste usually the responsibility of the generator.

There are a number of concepts about waste management which vary in their usage between countries or regions. Some of the most general, widely used concepts include:

249 

Waste

hierarchy:

The

waste

hierarchy

refers

to

the

"3

Rs" reduce, reuse and recycle, which classify waste management strategies according to their desirability in terms of waste minimization. The waste hierarchy remains the cornerstone of most waste minimization strategies. The aim of the waste hierarchy is to extract the maximum practical benefits from products and to generate the minimum amount of waste.



Extended producer responsibility: Extended Producer Responsibility (EPR) is a strategy designed to promote the integration of all costs associated with products throughout their life cycle (including end of life disposal costs) into the market price of the product. Extended producer responsibility is meant to impose accountability over the entire lifecycle of products and packaging introduced to the market. This means that firms which manufacture, import and/or sell products are required to be responsible for the products after their useful life as well as during manufacture.



Polluter pays principle: The Polluter Pays Principle is a principle where the polluting party pays for the impact caused to the environment. With respect to waste management, this generally refers to the requirement for a waste generator to pay for appropriate disposal of the waste.

6.1

Waste Management

Since the amount of gaseous produced by our plant is not too high, we have decided to release the gaseous to the atmosphere since it is not harmful to environment. The four main gas wastes are CO2, N2, H2 and O2 and all the gases are safe to be release. Few options have been taken into considerations such as recycling; reuse it or reducing the amount of these waste gasses release. Extra precaution and detailed planning is needed to successfully control and manage the waste from our plant as it is in the form of gas such as comparing the air of our plant to air quality standard.

250 6.2

Air –Quality Management Concepts

Air quality management is a term used to describe all the functions required to control the quality of the atmosphere (Anonymous, 1975). Among the essential elements of such a program are control regulations and a control strategy, legal authority to implement the control strategy, emission inventories, an atmospheric surveillance network, a data management system, agency staffing and funding, a system for analysis of complains and stack sampling production.

6.2.1 Air Quality Index

The Clean Air Act requires the EPA to establish National Ambient Air Quality Standards (NAAQS) for six pollutants considered harmful to public health and the environment: carbon monoxide, lead, nitrogen dioxide, ozone, particulate matter, and sulfur dioxide. The standards were set at the level required to provide an ample margin of safety to protect the public health (Peavy, 1985).

The Clean Air Act established two types of national air quality standards. Primary standards are intended to protect public health, including the health of "sensitive" populations such as asthmatics, children, and the elderly. Secondary standards set limits to protect public welfare, including protection against decreased visibility, damage to animals, crops, vegetation, and buildings.

251 National Ambient Air Quality Standards POLLUTANT

Carbon Monoxide (CO) 8-hour Average 1-hour Average

STANDARD VALUE

9 ppm 35 ppm

Lead (Pb) Quarterly Average Nitrogen Dioxide (NO2) Annual Arithmetic Mean Ozone (O3) 1-hour Average* 8-hour Average

STANDARD TYPE

10 mg/m3 Primary 40 mg/m3 Primary

1.5 µg/m3 Primary & Secondary

0.053 ppm

100 µg/m3 Primary & Secondary

0.12 ppm 0.08 ppm

235 µg/m3 Primary & Secondary 157 µg/m3 Primary & Secondary

Particulate < 10 micrometers (PM-10) Annual Arithmetic Mean 24-hour Average

50 µg/m3 Primary & Secondary 150 µg/m3 Primary & Secondary

Particulate < 2.5 micrometers (PM-2.5) Annual Arithmetic Mean 24-hour Average

15 µg/m3 Primary & Secondary 65 µg/m3 Primary & Secondary

Sulfur Dioxide (SO2) Annual Arithmetic Mean 24-hour Average 3-hour Average

0.03 ppm 0.14 ppm 0.50 ppm

80 µg/m3 Primary 365 µg/m3 Primary 1300 µg/m3 Secondary

* The ozone 1-hour standard applies only to areas that were designated nonattainment when the ozone 8-hour standard was adopted in July 1997. Figure 6.1: National Ambient Air Quality Standards (NAAQS)

252

CHAPTER 7

ECONOMIC ANALYSIS

7.1

Introduction

The economic analysis that involves the computation of the monetary investment  that  needs  to  poured  into  the  plant  should  comes   before,  because   it’s   one of the most essential parts of the entire process design. It also the determination of the relationship of income and expense to the material welfare of the company The calculation involves the economic potential for the selected plant was just a rough estimation in which the calculation does not involve in counting in other factors such as depreciation, plant lifetime and so on.

Economic evaluation of this project is done by estimating the fixed capital investment, total capital investment, total production cost and revenue from sales. Equipments costing have been calculated in Chapter V with the estimated Chemical Engineering Plant Index in 2009, which are 597.1. Most of the equipments quotations are obtained by employing the bare module method. Finally, the profitability analysis is performed by analyze the discounted cash flow. Payback period (PBP), discounted break-even period, net present value (NPV) and discounted cash flow rate of return (DCFRR) will be determined.

253

7.2

Grass-roof Capital

Equipment sizing has been carried out before and the cost for each of the equipment used in the plant is being estimated and because the gross root capital cost  is  the  major  portion  of  total  fixed  capital  cost,  it’s  has  been  done  first.  Cost  of   equipment, cost of installation, contingency and fees are all includes in the gross root capital cost part.

7.2.1

Equipment Costs and Grass Roof Capital

EQUIPMENT

COST ($)

Compressor

29 026.75

Compressor

27580.74

Compressor

157 559.35

Compressor

738 559.58

Compressor

187 101.75 890960.73 802244

Multi Tubular Fixed Bed Reactor Catalytic Tubular Fixed Bed Reactor Isomerazation Process Reactor

12552.41

Heat Exchanger

73500

Heat Exchanger

68250

Heat Exchanger Heat Exchanger Heat Exchanger Heat Exchanger

10080 106 122.8245 39480 135,000.00

Pump Pump

82469.84 61152

Pump

61152

Pump

61152

254 Pump Pump

94393.84 94393.84 33948.42 10 120.2709 782656.5445

Pump Pump Absorption Tower Distillation Column

726493.46

Distillation Column

893575

Distillation Column

2 474 200.972

Reboiler

193 585.4067

Reboiler

5460

Cooling Tower

17662.8

Storage Tank

551645.775

Storage Tank Total Bare Module Cost, Ctbm Total Bare Module Cost, Ctbm

157500 $ 9479580.5 RM 30457892.15

Table 7.1: Equipment Cost and Grass Root Capital

ITEMS

FORMULA

PRICE (RM)

Contingency and Fees

CC + CF = 0.08CTBM

2436631.37

Total Module Cost

CC + CF + CTBM

32894523.52

=CBM Auxiliary Facilities

0.10CTBM

Gross-roof Capital

3045789.22 38376944.1

(GRC) Table 7.2: Grass Root Capital

255

7.3

Fixed and Total Capital Investment Cost

Fixed Capital is the total cost for installed the process equipment with some auxiliaries that completed the operation of the process. It includes the cost of direct and indirect cost for the set up of the plant. Total Capital = Fixed Capital Investment + Working Capital + Start Up Cost Investment

Then, the working capital is the additional investment needed, in order to start up. The plant and operate to it point when income is earned. Working cost is most likely to be recovered at the end of the plant.

Specification

Cost (RM)

15%GRC

5,756,541.62

10%GRC

3,837,694.41

Piping (installed)

10%GRC

3,837,694.41

Electrical and material

5%GRC

1,918,847.21

Building

12%GRC

4,605,233.29

Yard improvements

3%GRC

1,151,308.32

Service facilities

8%GRC

3,070,155.53

Land

1%GRC

383,769.44

Onsite Purchased equipment installation Instrumentation and control (installed)

(installed)

Offsite

Total 1

Indirect cost

24,561,244.23

256 Engineering and

8%GRC

3,070,155.53

Legal expenses

1%GRC

383,769.44

Construction expenses

8%GRC

3,070,155.53

Contractor’s  fee

1.5%GRC

575,654.16

Contingency

5%GRC

1,918,847.21

supervision

Total 2

9,018,581.87

Total = total 1+ total 2

33,579,826.1

Gross root capital (GRC)

38,376,944.1

Fixed capital investment

71,956,770.2

(FCI) Working capital

10%FCI

7,195,677.02

Start up cost

5%FCI

3,597,838.51

Total capital investment

82,750,285.73

(TCI) Table 7.3: Fixed and Total Capital Investment Cost For the plant capital investment that have been calculated are includes all equipments cost, waste treatment cost, the direct cost in setting up the plant, indirect cost and also cost for working and start up cost. The total capital investment of RM 82,750,285.73 is needed where the calculations are tabulated in the Table 6.3.

7.4

Manufacturing Cost and Total Production Cost

Operating cost, the cost of producing the product with the estimation of raw materials cost, operating labor cost and also estimation of utilities cost for the plant.

257

7.4.1 Raw Materials Costing and Annual Profit

Raw material Butane

Price($/tone) 607.6

TOTAL ANNUAL COST Annual product credit 1-Butene 1,3-Butadiene Total

Amount(tonne/yr) 55082.984

Annual cost($/yr) 33,468,421.08 RM107,534,036.9

Cost

Amount

Annual cost

475.4 943.58

31407.09 50000

14,930,930.59 47,179,000 $ 62,109,930.59

TOTAL ANNUAL COST

RM 199,559,207

Table 7.4: Raw Materials Costing and Annual Profit

7.4.2 Estimation of Operating Labor Cost

Equipment type

Number of

Operator per Shift

Operator Required

equipment

per Equipment

Per Shift

Compressor

5

0.15

0.75

Reactor

3

0.5

1.5

Heat Exchanger

6

0.5

3

Pump

8

0

0

Absorption Tower

1

0.35

0.35

Distillation Column

3

0.35

1.05

Reboiler

2

0.5

1

Cooling Tower

1

0.35

0.35

Storage Tank

2

0

0

TOTAL

24

8

Table 7.5: Operating Labor Estimation

258 Operating labor = (4.5) (8) = 15.43 (rounding up to nearest integer yield 16 operators) = 36 operator/ days Labor cost in Malaysia = RM 1,500/months Therefore, Total labor cost = RM 54,000/month = RM 648,000/yr Direct supervision and clerical labor Assume that supervisory and clerical labor cost is 15% of operating labor cost. Cost of CDSCL = 0.15COL = 0.15 (RM 648,000) = RM 97,200.00/yr

7.4.3 Manufacturing Cost Summary

Fix capital investment (FCI): RM 71,956,770.2 Working capital: RM 7,195,677.02 Start up cost: RM 3,597,838.51 Total capital investment (TC1): RM 82,750,285.73 MANUFACTURING

SPECIFICATION

COST (RM)

EXPENSES Direct production cost Raw material: Butane

107,534,036.9

Utilities: Electricity

RM 0.26/kWh

848,474.40

Water

RM 1.20/m3

190,082.59

Steam

RM 14.25/1000 kg

9,845.22

Cooling water

RM 0.26/1000 kg

20,149.11

Maintenance and repairs

3%FCI

2,158,703.11

Operating supplies

0.5%FCI

359,783.85

Annual  employee’s  wages Patents and royalties

745,200 1%FCI

719,567.7

259

Indirect production cost Local taxes

1%FCI

719,567.7

Insurance

0.5%FCI

359,783.85

Plant overhead

50% L. Cost

35,978,380.1

Total manufacturing

114,037,794.4

expenses, AME General Expenses Administration Cost

5%FCI

3,597,838.01

distribution & selling

8%FCI

5,756,540.82

3%FCI

2,158,703.11

expenses Research & development Total general expenses,

11,513,081.94

AGE

Total production cost,

125,550,876.3

APC = AME + AGE(Excluding Depreciation) Depreciation, ABD

10%FCI

7,195,676.02

Total expenses, ATE

APC

132,746,552.4

Table 7.6: Total Expenses Estimation

Carried out calculation for rate of return, Revenue from sales, RM 1.05/kg Ethanol

= RM 199559207.00

Net annual profit, ANP

= RM 66812654.68

Income Taxes Of 30% Net Annual Profit = RM 20043796.404 Net Annual Profit After Income Taxes, ANNP = RM 46768858.276 Rate of return (ROR) = Cumulative net cash flow at end of project (100) Life of project X original investment

260

=

928497605.1593X 100 20 X 82750285.73

= 56.1%

7.5

Cash Flow Analysis

After all the calculation has been carried out, then the final step to be determined is the determination of payback period (PBP) and to solve that particular step, the value can be obtained by using graphical method. For the plant, the total operating period is 20 year with 3 year at the beginning as the start-up operation period. The capital investment used in the first year of the plant is  9  %  from  Total  Capital  Investment  (TCI)  while  it’s  increased  to  25  %  of  TCI.   In the third year of operation, capital investment to the plant account of 65 % of TCI plus the working capital.

To determine the time that must elapse after start up until cumulative undiscounted cash flow repays fixed capital investment, which is PBP, the cash flow diagram of undiscounted must be calculated. From the calculated value of net  cash  income  and  refer  to  cash  flow  diagram  that  has  been  plotted,  it’s  clearly that the Pay-back period (PBP) is estimated of 2 years after 2.5 years of start-up period.

261

Figure 7.1: Cash Flow Diagram with i= 0%

262

Figure 7.2: Discounted Cash Cumulative

Discounted break-even point is the point in which the time from the decision to proceed until discounted cumulative cash flow becomes positive value. And net present value is the final cumulative discounted cash flow value at project conclusion.

The discounted cash flow rate of return (DCFRR) is the point in which the rate result a Net present value of zero. In this case, the zero value cannot be obtained but the value is near the zero compared to other value, so the difference between the values can be neglected. So, the value of 40 % is selected as the break-even point. The DCFRR point is the maximum interest rate that counted after the taxes.

263 7.6

Conclusion

After carried out the economic analysis to the animal feed production plant, the conclusions that can be makes are:

 The Total Capital Investment is RM 82750285.73  The Total Production Cost is RM 125,550,876.3  The plant life is 20 years with the beginning 2.5 years as the start up period  From Discounted Cash flow diagram, Pay Back Period is estimated at 2 years after start up period.  The rate of returned obtained after the taxes is 56.1%.

264

CHAPTER 8

CONCLUSION AND RECOMMENDATION

8.0

Conclusion

As a conclusion, the proposed 50 000MT/Year of Butadiene from nbutane plant is considered to be very profitable plant. It is because only one existing plant of Butadiene in Malaysia with the cooperation of Titan Company. Thus, the competition for the product will be less and increase the potential of our Butadiene plant to expand the plastic and addition rubber market. This plant also capable to fulfill the demand of Butadiene and 1-butene in Malaysia for the next 20   years   because   in   2030,   it   is   estimated   that   about   2.5   millions   of   Malaysia’s   company want Butadiene production to produce another stuff like Butadiene rubber, ABS, SBR and so on.  At  Global  demand  for  butadiene  it’s  also  increase  at   an average annual rate of 3.9% during the period from 2001-2006 percent and will outpace capacity additions. This rate is higher than the compounded annual rate of 2.7% from 1996-2001 due to the global decline in demand that occurred in 2001 following the global economic slowdown

Besides, the Asian market has been also particularly active in building new capacity of butadiene and butadiene derivatives due to the ongoing development of automotive and tyre production in the region. The relocation of automotive industries increased synthetic rubber demand through tyre production.

265 So it is reliable to build Butadiene Plant for the overcome profit and cover for Malaysia, Asia Pacific and global demanded.

8.1

Recommendation

Based on the proposed plant, Butadiene also has water contamination from recycling process in the plant. Hence, several strategies need to be sort out to overcome this problem. Good waste management need to be carried out in order to protect the environment which is build a small pond of waste water to be treated in the proposed plant.

Other than that, the energy consumption and manufacturing cost to build this proposed plant is quite high. In order to minimize the cost, we can reduce it with develop heat integration using Pinch technology, otherwise to increasing production rates for maximizing profit. Besides the safety aspect of the plant can be upgraded

266

REFERENCE

1. Butadiene production process overview;Chemico-Biological Interactions; Volume 166, Issues 1-3, 20 March 2007, Pages 10-14,Wm. Claude White. 2. ‘Ullmann’s  Encyclopedia  of  Industrial  Chemistry  5th  completely  revised   edition  Vol.  4’,Butadiene, VCH Publishers. 3. ‘Kirk-Othmer Encyclopedia of Chemical Technology  5th  edition  Vol.4’,   Butadiene, John Wiley & Sons Inc. Publisher. 4. A  Chauvel,  G.Lefebvre,  ‘Petrochemical  processes-synthesis gas derivatives  and  major  hydrocarbons’,  Editions  Technip,  Paris  1989 5. An  Ullmann’s  encyclopedia,’Industrial  organic  chemicals-starting materials  and  intermediates,  volume  2’,  Wiley~VCH,  1999 6. Albert.V.G.Hahn,  ‘The  petrochemical  industry-Market  and  economics’,   McGraw Hill Book Company, 1970 7. Robert  H.Perry,  Don  W.Green,  ‘Perry’s  chemical  engineers’  handbook’,   7th edition, McGraw Hill Company, 1997 8. David  R.Lide(chief  editor),  ‘CRC  handbook  of  chemistry  and  physics’,   86th edition, CRC Press, Taylor & Francis Group, LLC, 2005. 9. 1,3-Butadiene: http://en.wikipedia.org/wiki/Butadiene (accessed 16 August 2009) 10. Titan inks RM536m sales deal with Kumho http://www.mida.gov.my/beta/news/print_news.php?id=2111 (accessed 16 August 2009) 11. CMAI Completes 2004 World Butadiene Analysis http://www.prnewswire.com/cgibin/stories.pl?ACCT=104&STORY=/www/story/12-012003/0002066964&EDATE (accessed 6 August 2009) 12. 28-08-2005: Titan Chemicals secures RM500m contract http://www.theedgedaily.com/cms/contentPrint.jsp?id=com.tms.cms.articl e.Article_f16c6 cf5-cb73c03a-c5577c00-a6616fec&paging=0# (accessed 11 August 2009)

267

13. Butadiene http://www.chemlocus.com/news/newssub_read.htm?menu=D17&Sequen ce=6909&cpag e=1 (accessed 6 August 2009) 14. Shell Chemicals – Butadiene http://www.shellchemicals.com/butadiene/1,1098,1494,00.html (accessed 16 August 2009) 15. Butadiene – Product Stewardship Guidance Manual http://www.shellchemicals.com/chemicals/pdf/butadiene/butadiene_ps_m an10_02.pdf?se ction=our_products (accessed 1 November 2009) 16. CMAI Completes 2002 World Butadiene Analysis http://www.cmaiglobal.com/news/WBA2002.pdf (accessed 25 Jun 2009) 17. CMAI Announces Completion of Olefins World Analyses; World Light Olefins Analysis and World Butadiene Analysis http://www.redorbit.com/news/science/755600/cmai_announces_completi on_of_olefins_ world_analyses_world_light_olefins/index.html?source=r_science (accessed 1 July 2009) 18. Method for the production of butadiene from n-butane http://www.freepatentsonline.com/7034195.html (accessed 16 July 2009) 19. Butadiene production and purification http://www.freepatentsonline.com/4054613.html (accessed 16 July 2009) 20. Butadiene recovery process http://www.freepatentsonline.com/4038156.html (accessed 16 July 2009) 21. Method for obtaining crude 1,3-butadiene from a c4 fraction http://www.freshpatents.com/Method-for-obtaining-crude-13-butadienefrom-a-c4- fraction-dt20070222ptan20070039813.php (accessed 16 January 2010) 22. Butadiene Extraction Technology http://www02.abb.com/GLOBAL/NOOFS/noofs187.nsf/viewunid/2B719 2CD2C0193A4 85256BDD006FAD27/$file/Butadiene+2Web.pdf (accessed 16 January 2010) 23. Butadiene production http://www.sciencedirect.com/science?_ob=ArticleURL&_udi=B6T564MX56J31&_user=708426&_coverDate=03%2F20%2F2007&_alid=600557633&_ rdoc=2&_fmt= full&_orig=search&_cdi=4994&_sort=d&_docanchor=&view=c&_ct=6& _acct=C00001

268 2438&_version=1&_urlVersion=0&_userid=708426&md5=ae0a35aebf8b c435711b3f86 2bda5533#secx3 (accessed 16 January 2010) 24. Process Economics Program Report 35C http://www.sriconsulting.com/PEP/Public/Reports/Phase_94/RP035C/ (accessed 31 January 2010) 25. Process Economics Program Report 35B http://www.sriconsulting.com/PEP/Public/Reports/Phase_80/RP035B/ (accessed 19 February 2010) 26. Introduction to Process Technology http://www.dmtcalaska.org/course_dev/processtech/04/notes04.html (accessed 6 February 2010) 27. Process for producing butadiene http://www.freepatentsonline.com/4595788.html (accessed 1 March 2010) 28. Material Safety Data Sheet (n-butane) http://www.marathonpetroleum.com/MSDS/0105MAR019.pdf (accessed 1 March 2010) 29. Material Safety Data Sheet (1-butene) http://www.mathesontrigas.com/pdfs/msds/MAT25870.pdf (accessed 6 March 2010) 30. Safety data for 1,3-butadiene http://ptcl.chem.ox.ac.uk/MSDS/BU/1,3-butadiene.html (accessed 6 March 2010) 31. Process Plant and Machinery http://www.processplantandmachinery.com/ (accessed 6 March 2010) 32. Chemical Process Equipment http://www.chemicalplant.com (accessed16 March 2010) 33. Pump http://en.wikipedia.org/wiki/pump (accessed 16 March 2010) 34. Construction Equipment http://www.constructionequiptment.com/ (accessed 16 March 2010) 35. Heat Exchanger http://www.engineersedge.com/heat_exchanger (accessed 16 March 2010) 36. Products and Supplier http://material-properties.globalspec.com (accessed 4 April 2010) 37. Solvent Stripper

269 http://products.21food.com/a-Solvent%20Stripper (accessed 4 April 2010) 38. Butadiene http://www.sriconsulting.com/CEH/Public/Reports/444.0000/ (accessed 4 April 2010) 39. Toray to Lift ABS Resin Output Capacity at Malaysia Unit http://www.redorbit.com/news/science/474278/toray_to_lift_abs_resin_ou tput_capacity_at_malaysia_unit/?source=r_science (accessed 4 April 2010) 40. Malaysia - Titan net profit surges 154 percent in Q3 http://www.yarnsandfibers.com/news/index_fullstory.php3?id=10466 (accessed 4 April 2010) 41. Material Safety Data Sheet (MOPN) http://www.dixiechemical.com/pdf/3 Methoxypropionitrile%20(MOPN)%20DIXIE%20MSDS.pdf (accessed 2 January 2010) 42. Richard M. Felder, Ronald W. Rousseau,  ‘Elementary  Principles  of   Chemical  Processes’,  3rd  Edition,  John  Wiley  &  Sons,  Inc. 43. J.  M.  Smith,  H.C.  Van  Ness,  M.  M.  Abbott,  ‘Introduction  to  Chemical   Engineering  Thermodynamics’,  6th  edition,  McGraw-Hill International 44. Max S. Peters, Klaus D. Timmerhaus,  Ronald  E.  West,  ‘Plant  Design  and   Economics  for  Chemical  Engineers’,  5th  edition,  McGraw-Hill International. 45. R.K.  Sinnott,  ‘Coulson  &  Richardson’s  Chemical  Engineering  Design,   Volume  6’,  3rd edition, Butterworth Heinemann. 46. Stanley  M.  Walas,  ‘Chemical  Process  Equipment  Selection  and  Design’,   Butterworth Heinemann, Boston, 1990. 47. Anonymous, (February 2009),Butadiene Uses and Market Data, Website : http://www.icis.com/v2/chemicals/9075172/butadiene/uses. 48. Anonymous, (February 2009),Butadiene CAS No: 106-99-0, Website : http://www.icis.com/v2/chemicals/9075170/butadiene

270 49. Anonymous, (February 2009),HistoryButadiene, Website : http://en.wikipedia.org/wiki/1,3-Butadiene

50. Armstrong, H.E. Miller, A.K. (1886). "The decomposition and genesis of hydrocarbons at high temperatures. I. The products of the manufacture of gas from petroleum." Journal of the Chemical Society 49, 80. 51. Sun, H.P. Wristers, J.P. (1992). Butadiene. In J.I. Kroschwitz (Ed.), Encyclopedia of Chemical Technology, 4th ed., vol. 4, pp. 663–690. New York: John Wiley & Sons.

52. Beychok, M.R. and Brack, W.J., "First Postwar Butadiene Plant", Petroleum Refiner, June 1957.

53. Kirshenbaum, I. (1978). Butadiene. In M. Grayson (Ed.), Encyclopedia of Chemical Technology, 3rd ed., vol. 4, pp. 313–337. New York: John Wiley & Sons.

54. Anonymous, (June 21, 2007),Product Safety Assessment, Butadiene, Website : http://www.dow.com/PublishedLiterature/dh_007e/0901b8038007ee62.pd f?filepath=productsafety/pdfs/noreg/233-00281.pdf&fromPage=GetDo

55. Anonymous, (1987), Synthetic elastomeric with improved chemical, aging and oil resistance, Website : http://www.freepatentsonline.com/4789708.html

56. Anonymous, (2007), Engineering Resin, Website : http://www.alliancepoly.com/engineering.asp

57. Anonymous, (2007), Types of Synthetic Rubber, Website : http://orientalchemic.com/rubbers.html

271 58. Anonymous, (February 2009),Butadiene Uses and Market Data, Website : http://www.icis.com/v2/chemicals/9075172/butadiene/uses

59. Anonymous, (November 1996),Locating and Estimating Air Emissons From Source, Website : http://www.scribd.com/doc/1829848/Environmental-Protection-Agencybutadien 60. Dr. R. Talishinsky, (1996),Butadiene production, Website : http://www.eldanel.com/docs/buta.pdf

61. Lynne M.Miller, (Dec 1978),Investigate of Selected Potential Environmental Contaminants : Butadiene and its Oligomers, Website : http://www.ncbi.nlm.nih.gov/pubmed/17324391

62. Anonymous, (January 2010),Butadiene, Website : http://www.sriconsulting.com/WP/Public/Reports/butadiene/ 63. Jorg Wutke, (1996),The petrochemical Industry in China, Website : http://www.britannica.com/EBchecked/topic/86395/butadiene#ref239145

64. Anonymous, (January,14,2008), CMAI Completes 2008 World Butadiene Analysis, Website :http://www.cmaiglobal.com/Marketing/News/WBA2008.pdf

65. Anonymous, (2008), Butadiene Derivatives Impacted by Automotive Crisis Website : http://www.chemsystems.com/about/cs/news/items/PPE%20PCMD%20B utadiene%20Derivatives%202009.cfm 66. Anonymous, (1996),Butadiene rubber, Website : http://www.eldanel.com/docs/buta.pdf

272

67. Anonymous, (November 1996),Butadiene Styrene, Website : http://www.industrialrubbergoods.com/styrene-butadiene-rubber.html

68. Anonymous, (2007),Product overview and market projection of emerging bio , Website: http://www.criterioncatalysts.com/home/content/chemicals/products_servi ces/our_products/lower_olefins/butadiene/product_overview/butadiene_o verview.html

69. Anonymous, (2009),Butadiene Market Dynamics, Website:: http://www.chemsystems.com/about/cs/news/items/PPE%20Butadiene%2 0MarketDynamics.cfm

70. Anonymous, (2008),Basic Material: Global Insights, Website : http://www.fullermoney.com/content/2005-0112/RameshMorganSmetalsms2004dec14global.pdf

71. Anonymous, (2008),Management Discussions, Website : http://www.indiainfoline.com/Markets/Company/Fundamentals/Managem ent-Discussions/Reliance-Industries-Ltd/500325

72. Anonymous, (Nov 2008),Production from China, Website : http://www.reuters.com/article/pressRelease/idUS120873+01-Dec2008+PRN20081201

73. Anonymous, (2002),Butadiene product Stewardship Guidance Manual , Website : http://wwwstatic.shell.com/static/chemicals/downloads/products_services/ butadiene_ps_man10_02.pdf

273 74. Anonymous, (November 1996),Produce of Butadiene, Website : http://digitalibrary.mida.gov.my/equipmida/custom/indReports/petrochem ical_polymer/2008/PetrochemMal2008.pdf

75. Anonymous, (November 2006),Production of Butadiene, Website : http://www.purchasing.com/article/220145Butadiene_prices_on_the_way_down_to_30_lb.php 76. Anonymous, (2007),Butadiene Market Demand, Website: http://www.forbes.com/feeds/businesswire/2009/08/20/businesswire12807 6278.html 77. Anonymous, (1996),Process of Butadiene production, Website :http://www.plastemart.com/plasticnews_desc.asp?news_id=9421 78. Peters,   M.S.,   Timmerhaus,   K.D.,   and   West,   R.E.,   (2004),   “Plant   Design   and  Economics  for  Chemical  Engineers”,  New  York,  Mc  Graw  Hill 79. Anonymous, (1975), Air Quality Management, in Air Pollution Control Orientation Course (422-A), U.S. EPA, Atlanta 80. Anonymous,

(2009),

Waste

Management,

Website:

http://en.wikipedia.org/wiki/Waste_management#cite_noteWaste_Management_FAQ-0 81. Peavy,  H.S.,  Rowe,  D.R.,  and  Tchobanoglous,  G.,  (1985),  “Environmental   Engineering,”  New  York,  Mc  Graw  Hill  International Editions

274

APPENDIX A

MANUAL CALCULATION OF ENERGY BALANCE

275 Manual Calculation of Energy Balance Heat Capacity CP = A + BT + CT2 + DT3 Components Butane 1-Butene 2-butene 1,3-Butadiene Water Carbon dioxide Hydrogen Oxygen Nitrogen

A 92.30x10-3

B 27.88x10-5

C -15.47x10-8

D 34.98x10-12

33.46x10-3 36.11 x10-3

0.6880x10-5 4.233 x10-5

0.7604x10-8 -2.887 x10-8

28.84 x10-3

0.00765 x10-5

29.10 x10-3

1.158 x10-5

0.3288 x10-8 -0.6076 x10-

-3.593x10-12 7.464 x10-12 -0.8698 x10-

29.00 x10-3

0.2199 x10-5

Desired energy

=

x (Ht + Hf)

Desired energy

=

x

=

8

0.5723 x10-8

12

1.311 x10-12 -2.871 x1012

+P

Stream 1 N-Butane desired energy

= 118.71 mole/hr x (7540 kJ/mole -124.7 kJ/mole) = 880270.263 kJ/hr

Total desired energy = 880270.263.kJ/hr Simulation value for stream 1 = 0.196MMkcal/hr Percentage error of energy is Simulation = 0.195x 106 kcal/hr x 4.1858j/cal = 816231kJ/hr Percentage error =

x 100%

276

Percentage error =

x 100% = 7.27%

Stream 2 Hydrogen: desired energy

= 1077.78 mole/hr x (550.18 kJ/mole) = 593641.224 kJ/hr

Oxygen desired energy

= 134.49 mole/hr x (88028.8 kJ/mole) = 11838993.31 kJ/hr

Nitrogen desired energy

= 308.04 mole/hr x (77025.2kJ/mole) = 23726842.61 kJ/hr

Total desired energy = 593641.224 kJ/hr + 11838993.31 kJ/hr + 23726842.61 kJ/hr = 36159477.14 kJ/hr Simulation value for stream 2 = 8.196MMkcal/hr Percentage error of energy is Simulation = 8.235x 106 kcal/hr x 4.1858j/cal = 34470063kJ/hr Percentage error =

x 100%

Percentage error =

x 100% = 4.672%

Stream 3b Water: desired energy

= 296.76 mole/hr x (50238 kJ/mole-241.83 kJ/mole) = 14836863.41 kJ/hr

Total desired energy = 14836863.41 kJ/hr Simulation value for stream 3b = 8.783MMkcal/hr

277 Percentage error of energy is Simulation = 3.335x 106 kcal/hr x 4.1858j/cal = 13959643kJ/hr Percentage error =

x 100%

Percentage error =

x 100% = 5.91%

Stream 5 N-butane: desired energy

= 118.7133 mole/hr x (62884.548 kJ/mole-124.7 kJ/mole) = 7450428.664 kJ/hr

1-butene: desired energy

= 40.52161 mole/hr x (54976.598 kJ/mole + 1.17 kJ/mole) = 2227787.674 kJ/hr

2-butene: desired energy

= 86.9925 mole/hr x (54709.397kJ/mole + 1.17 kJ/mole) = 4759409 kJ/hr

1,3 butadiene: desired energy

= 8.706111mole/hr x (50283.658kJ/mole -956 kJ/mole) = 429452.0166 kJ/hr

Water: desired energy

= 296.7622 mole/hr x (15107.457 kJ/mole – 241.83kJ/mole) = 4411556.173 kJ/hr

Carbon dioxide: desired energy

= 5.468636 mole/hr x (18599.423 kJ/mole – 393.5 kJ/mole) = 99561.5659 kJ/hr

Hydrogen:

278

desired energy

= 259.7053 mole/hr x (12328.319 kJ/mole ) = 3201729.784 kJ/hr

Nitrogen; desired energy

= 4312.615 mole/hr x (12849.743kJ/mole) = 55415994.41kJ/hr

Total desired energy = 7450428.664 kJ/hr + 2227787.674 kJ/hr + 4759409 kJ/hr + 429452.0166 kJ/hr + 4411556.173 kJ/hr + 99561.5659 kJ/hr + 3201729.784 kJ/hr + 55415994.41kJ/hr = 77995919.29 kJ/hr Simulation value for stream 3b = 8.783MMkcal/hr Percentage error of energy is Simulation = 17.395x 106 kcal/hr x 4.1858j/cal = 72811991kJ/hr Percentage error =

x 100%

Percentage error =

x 100% = 6.65%

Stream 7 N-butane: desired energy

= 118.7133mole/hr x (62884.548 kJ/mole-147.0 kJ/mole) = 559606.1638 kJ/hr

2-butene: desired energy

= 55.95121mole/hr x (54709.397kJ/mole + 1.17 kJ/mole) = 3061122.423 kJ/hr

1,3 butadiene: desired energy

= 79.4958mole/hr x (50283.658kJ/mole - 1127 kJ/mole) = 3907747.853 kJ/hr

279 Water: desired energy

= 544.2868mole/hr x (15107.457kJ/mole -285.84 kJ/mole) = 8067210.488 kJ/hr

Carbon dioxide: desired energy

= 20.1344mole/hr x (18599.423kJ/mole – 412.9kJ/mole) = 366174.7287 kJ/hr

Hydrogen: desired energy

= 81.42588 mole/hr x (12328.319kJ/mole-0) = 1003844.224 kJ/hr

Oxygen: desired energy

= 1044.578 mole/hr x (13249.802kJ/mole ) = 13840451.67 kJ/hr

Nitrogen; desired energy

= 4146.206mole/hr x (12849.743kJ/mole) = 53277681.53kJ/hr

Total desired energy = 559606.1638 kJ/hr + 3061122.423 kJ/hr + 3907747.853 kJ/hr + 8067210.488 kJ/hr + 366174.7287 kJ/hr + 1003844.224 kJ/hr + 13840451.67 kJ/hr + 53277681.53kJ/hr = 84083839.08 kJ/hr Simulation value for stream 3b = 8.783MMkcal/hr Percentage error of energy is Simulation = 18.365x 106 kcal/hr x 4.1858j/cal = 76872217kJ/hr Percentage error = Percentage error =

x 100% x 100% = 8.57%

280 Stream 9a N-butane: desired energy

= 172.9575845 mole/hr x (2122.514kJ/mole-147.0 kJ/mole) = 341680.1296 kJ/hr

2-butene: desired energy

= 63.7439498mole/hr x (7189.784kJ/mole + 1.17 kJ/mole) = 458379.8108 kJ/hr

1,3 butadiene: desired energy

= 116.0073471mole/hr x (6609.026 kJ/mole - 1127kJ/mole) = 635955.293 kJ/hr

Water: desired energy

= 40.68929584 mole/hr x (2543.847kJ/mole -285.84

kJ/mole) = 91876.71483 kJ/hr Carbon dioxide: desired energy

= 28.91870004mole/hr x (3038.834kJ/mole – 412.9kJ/mole) = 75938.5977 kJ/hr

Hydrogen: desired energy

= 6005.393935 mole/hr x (2122.514 kJ/mole – 0) = 12746532.7 kJ/hr

Oxygen: desired energy

= 119.3204777mole/hr x (2264.693kJ/mole ) = 270224.2506 kJ/hr

Nitrogen; desired energy

= 1524.877843 mole/hr x (2197.475/mole) = 3350880.938 kJ/hr

Total desired energy = 341680.1296 kJ/hr + 458379.8108 kJ/hr + 635955.293 kJ/hr +

281 91876.71483 kJ/hr + 75938.5977 kJ/hr + 12746532.7 kJ/hr + 270224.2506 kJ/hr + 3350880.938 kJ/hr = 17399468.43 kJ/hr Simulation value for stream 9a = 4.0MMkcal/hr Percentage error of energy is Simulation = 4.0x 106 kcal/hr x 4.1858j/cal = 72811991kJ/hr Percentage error =

x 100%

Percentage error =

x 100% = 3.77%

Stream 11 N-butane: desired energy

= 172.3704507mole/hr x (1491.535kJ/mole-147.0 kJ/mole) = 231758.1039 kJ/hr

2-butene: desired energy

= 62.95073242mole/hr x (1340.899kJ/mole + 1.17 kJ/mole) = 84484.2265 kJ/hr

1,3 butadiene: desired energy

= 115.5460243mole/hr x (1220.147kJ/mole - 1127kJ/mole) = 10762.7655 kJ/hr

Water: desired energy

= 31.23550603 mole/hr x (1125.511kJ/mole -285.84 kJ/mole) = 26227.5486 kJ/hr

Carbon dioxide: desired energy

= 8.445301996 mole/hr x (598.945 kJ/mole – 412.9kJ/mole) = 1571.20621 kJ/hr

Oxygen:

282

desired energy

= 0.584009114mole/hr x (450.326kJ/mole – 0) = 262.9945 kJ/hr

Nitrogen; desired energy

= 119.3204777mole/hr x (436.974kJ/mole ) = 52139.9462 kJ/hr

Total desired energy = 231758.1039 kJ/hr + 84484.2265 kJ/hr + 10762.7655 kJ/hr + 10762.7655 kJ/hr + 26227.5486 kJ/hr + 1571.20621 kJ/hr + 262.9945 kJ/hr + 52139.9462 kJ/hr = 417969.5568 kJ/hr

Simulation value for stream 9a = 2.457MMkcal/hr Percentage error of energy is Simulation = 0.095x 106 kcal/hr x 4.1858j/cal = 397651kJ/hr Percentage error =

x 100%

Percentage error =

x 100% =4.8%

Stream 13 N-butane: desired energy

= 1.0896mole/hr x (209.3kJ/mole-147.0 kJ/mole) = 67.888kJ/hr

1,3 butadiene: desired energy

= 115.8566mole/hr x (7251.714kJ/mole - 1127kJ/mole) = 709588.54kJ/hr

Total desired energy = 67.888kJ/hr + 709588.54kJ/hr = 709656.428 kJ/hr

283 Simulation value for stream 9a = 4.0MMkcal/hr Percentage error of energy is Simulation = 0.15x 106 kcal/hr x 4.1858j/cal = 627870 kJ/hr Percentage error =

x 100%

Percentage error =

x 100% = 11.52%

Stream 15 N-butane: desired energy

= 6.965058584 mole/hr x (887.404kJ/mole-147.0 kJ/mole) = 5156.9572 kJ/hr

1-butene: desired energy

= 110.3348351mole/hr x (780.367kJ/mole + 1.17 kJ/mole) = 86230.7559 kJ/hr

Carbon dioxide desired energy

= 11.36487036mole/hr x (358.833kJ/mole - 412.9 kJ/mole) = -614.4644 kJ/hr

Oxygen: desired energy

= 0.422452344mole/hr x (270.038kJ/mole) = 114.0782 kJ/hr

Nitrogen desired energy

= 0.177786569mole/hr x (262.032kJ/mole ) = 46.5858 kJ/hr

Total desired energy = 5156.9572 kJ/hr + 86230.7559 kJ/hr -614.4644 kJ/hr + 114.0782 kJ/hr + 46.5858 kJ/hr = 90933.9127 kJ/hr Simulation value for stream 15 = 8.783MMkcal/hr

284 Percentage error of energy is Simulation = 0.02x 106 kcal/hr x 4.1858j/cal = 83716kJ/hr Percentage error = Percentage error =

x 100% x 100% = 7.94%

285

APPENDIX B

EQUIPMENT SIZING

286 Reactor specification Identification number : R-102 Number required : 4 ( parallel) Function General Type of reactor orientation operation Feed (kmol/h) nbutane 1butene c2butene 1,3butadiene steam carbon dioxide hydrogen nitrogen oxygen Total flow Temperature Pressure Catalyst Type Volume Weight Shape Conversion of C4mixture : selectivity of 1,3 butadiene yield holding method Tube Number length catalyst fill height Inside diameter Outside diameter Gauge Heat transfer area passes Tube arrangement pattern Tube pitch Pressure drop Inlet temperature Outlet temperature Shell Fluid fluid composition Flow rate

Date : 2/3/2010 By : Kumaran To convert n-butane feed to1,3butadiene catalytic tubular fixed bed vertical continuous 246.3758717 84.11562543 180.5810164 18.0764008 616.6877427 11.37091046 8573.518106 0 640.3678968 10371.09 450(°C) 2.0 atm Bismuth molybdate(10wt% αBiMo3O12 and 90% γ  BiMo2O  12) 422.72 m3 608716.8 kg 6.0 mm spherical pellets 0.6 2mol C4H6/1molCO2 formed 0.6 wire gauze and clamping ring 34760 6.0 m 6.0m 25.4 mm 31.8mm BWG 12(heavy) 7882m2 1(up flow) square 39.7 mm 1atm 450°C 500°C Hitec heat transfer salt 40% NaNO2. 7%NaNo3, 53%KNO3 10371.09kmol/h

287 Inlet temperature outlet temperature Heat removed per tube Heat transfer coefficient Inside diameter Passes Baffles Pressure drop Materials of construction ASTM code General description Grade and class Composition tensile strength Yield strength Mechanical design of reactor vessel Overall height outside diameter Total weight(empty) Design pressure MAWP wall thickness Tube plate Type of ends ellipsoidal, end thickness Gas connection

Insulation Type Thermal conductivity Thickness

144oC 600oC 253 W 40 W/m2K 6.49m 1 5(segmental) 242.31Pa A387 : Pressure vessel plates, alloy steel, chromium molybdenum High alloy chromium steel stainless steel Type 316 18% Cr, 8% Ni, 2.5% Mo, 500MPa-650MPa 310 MPa 8.0 m 6.49m 240 T 22 atm 22 atm 20mm 25mm 2.5:1 21mm 1492mm steel pipe

Mineral fibre double blanket 0.036 W/mK 150 mm(2x75mm)

288 DISTILLATION COLUMN Identification: Item: Item Number: D-103 No. Required: 1

Date: 1-Mac-10 By: Wan M. Irman

Function:To recover n-butane and 1-butene Operation: Continuous Materials handled: Feed Bottoms Quantity (kg/hr) 7399.3551 Component (kg/hr) n-butane 899.899 1-butene 2-butene 461.719 Water 36.997 Carbon Dioxide Oxygen Nitrogen Temperature, °C

Overhead

Reflux

7117.66732

34323.527

7791.010

404.995

1953.007

6577.008 325.594

6190.235 -

15504.4987

434.126 356.603

116

29851.17

-

499.66

2409.51

13.954

13.524

65.217

6.202

9.253

44.621

115

144 Design Data: No. of trays: 31 Feed stage (from bottom):3 Pressure, atm: Top of the column: 3.5 Bottom of the column: 4.5 Reflux ratio: 4.82 Functional height, m: 31.40 Tray spacing, m: 0.91 Material of construction: Stainless Steel 316 Liquid density, kmol/m : 8.23 Vapor density, kmol/m : 0.11 Maximum allowable vapor velocity (superficial), m/s: 0.64 Maximum vapor flow rate, m /s: 760.06 Recommended inside diameter, m: 2.13 Tray efficincy, %: 60 Utilities: Cooling water, steam Control: Reflux ratio, column pressure, number of trays

115

289 Tolerances: Heuristic Comment and Drawing: Included

CONDENSER Identification: Item: Date : 1-Mac-10 Item Number: E-107 By : Wan M. Irman No. Required: 1 Function: To condense the overhead vapor from distillation column D-103 Operation: Continuous Type: Shell and Tube Fixed-tube Tube side: Fluid handled: cooling water Flow rate: 2062.809 kg/hr Temperature: 30°C to 45°C Pressure: 1 atm Pressure drop: 0.27 atm Material of construction: Carbon steel Shell Side: Fluid handled: Overhead vapors mm Flow rate: 11.86 kg/s mm Temperature: 115°C Pressure: 3.5 atm Pressure drop: 0.019 atm Material of construction: Carbon steel Overall Coefficient: 553.16 W/m°C

Tubes: Outside diameter: 50 mm Inside diameter: 46 mm Length: 610 mm Pitch: Arrangement Passes: 4 Tube per pass: 86 Shell: Baffle spacing: 1502.81 Inside diameter: 1502.81 Passes: 1

Utilities: Cooling water Control: Tubular Exchangers Manufacturers Association (TEMA) Standards Tolerances: Heuristic Comment and Drawing: Included

290 REBOILER Identification: Item: Date : 1-Mac-10 Item Number: E-108 By : Wan M. Irman No. Required: 1 Function: To vaporize partial of the bottom product from distillation column D103 Operation: Continuous Type: Shell and Tube U-tube Tube side: Fluid handled: Dowtherm A Flow rate: 72108 kg/hr Temperature: 305°C mm Pressure: 1 atm Material of construction: Carbon steel Shell Side: Fluid handled: Overhead vapors mm Flow rate: 7117.66732 kg/hr Temperature: 115°C Pressure: 0.019 atm Material of construction: Carbon steel Overall Coefficient: 553.16 W/m°C

Tubes: Outside diameter: 30 mm Inside diameter: 25 mm Nominal Length: 10400 Pitch: Arrangement Passes: 2 Tube per pass: 340 Shell: Inside diameter: 1502.81 Passes: 1

Utilities: Cooling water Control: Tubular Exchangers Manufacturers Association (TEMA) Standards Tolerances: Heuristic Comment and Drawing: Included

291 PUMP Identification Function

Item Type : Centrifugal pump Date : 1 Mac 10 Item No : P-105A/B By : Wan M. Irman No Required : 1 Used to move reactant or liquids from one place to another by mechanical means (lower pressure to higher pressure).

Operation

Continuous

Orientation & Dimension

Mass Flow Rate : 7117.66732 kg/h Shaft Power : 0.697 kW Fluid Power : : 0.557 kW

Operating Condition

Pressure Difference : 1 atm Temperature : 34°C

Material of Construction

-Alloys including Alloy 20, 317L and Hastelloy. -Low temperature materials

Utilities

The electricity for power requirement is bought from TNB

Controls

The volumetric flowrate and pressure required.

Tolerances

Based on reference and heuristics rules Ref  ,  Mc  Cabe  Smith  Harriot,’Unit  Operation  of  chemical engineering’Sixth  Edition)

292 PUMP Identification

Function

Item Type : Centrifugal pump Date : 1 Mac 10 Item No : P-106A/B By : Wan M. Irman No Required : 1 Used to move reactant or liquids from one place to another by mechanical means (lower pressure to higher pressure).

Operation

Continuous

Orientation & Dimension

Mass Flow Rate :7117.66732 kg/h Shaft Power : 0.697 kW Fluid Power : 0.557 kW

Operating Condition

Pressure Difference : 1atm Temperature : 34°C

Material of Construction

-Alloys including Alloy 20, 317L and Hastelloy. -Low temperature materials

Utilities

The electricity for power requirement is bought from TNB

Controls

The volumetric flowrate and pressure required.

Tolerances

Based on reference and heuristics rules Ref  ,  Mc  Cabe  Smith  Harriot,’Unit  Operation  of  chemical engineering’Sixth  Edition)

293 PUMP Identification Function

Item Type : Centrifugal pump Date : 1 Mac 10 Item No : P-107A/B By : Wan M. Irman No Required : 1 Used to move reactant or liquids from one place to another by mechanical means (lower pressure to higher pressure).

Operation

Continuous

Orientation & Dimension

Mass Flow Rate :7399.3551 kg/h Shaft Power : 0.362 kW Fluid Power : 0.289 kW

Operating Condition

Pressure Difference : 0.5 atm Temperature : 115°C

Material of Construction

-Alloys including Alloy 20, 317L and Hastelloy. -Low temperature materials

Utilities

The electricity for power requirement is bought from TNB

Controls

The volumetric flowrate and pressure required.

Tolerances

Based on reference and heuristics rules Ref  ,  Mc  Cabe  Smith  Harriot,’Unit  Operation  of  chemical engineering’Sixth  Edition)

294 COMPRESSOR Identification : Function : Operation : Type : Materials Handled (kmol/hr)

Design Data

Item Compressor Date : 1 Feb 2010 Item No. C-106A/B No. required 1 By : Kumaran Use to increase the pressure on a fluid to transport it through a pipe. Continuous Flow Centrifugal Single Stage Compressor CO2=1293.42882 H2=1430.380107 O2=3804.202412 N2=42702.17207

Inlet pressure : 1 atm Outlet pressure : 2 atm Inlet temperature : 450 °C Outlet temperature : 470°C Pressure  change,  ΔP                      : 1 atm Compressor power : 1462.84kW

Utilities :

Compressor efficiency No. of stages Compressor ratio Electricity supply

Material of Construction : Tolerances :

Stainless steel Rules of thumb for compressor

Storage Tank

: 80 % : 1 (single stage) :4

TK-2

Identification: storage tank Item No: TK-2 No required: 1 Function: to store the 1-Butene Operation: hold tank Type: Vertical tank on concrete pad Temp: 34 ºC Design Data Diameter: 2.33 m Height: 6.978 m Volume: 29.656 m Material of construction: Stainless steel type 302 Insulation: Not provided Comments and Drawings: Not provided

By: Wan M. Irman Date: 1-Mac-10

295

CONDENSER Identification: Item: Item Number: E-106 No. Required: 1

Date : 1-Mac-10 By : Wan M. Irman

Function: To condense the overhead vapor from distillation column D-103 Operation: Continuous Type: Shell and Tube Fixed-tube Tube side: Fluid handled: cooling water mm Flow rate: 2062.809 kg/hr Temperature: 30°C to 45°C Pressure: 1 atm Pressure drop: 0.27 atm Material of construction: Carbon steel

Tubes: Outside diameter: 50

Shell Side: Fluid handled: Overhead vapors mm Flow rate: 11.86 kg/s 1502.81 mm Temperature: 115°C Pressure: 3.5 atm Pressure drop: 0.019 atm Material of construction: Carbon steel Overall Coefficient: 553.16 W/m°C

Shell: Baffle spacing: 1502.81

Inside diameter: 46 mm Length: 610 mm Pitch: Arrangement Passes: 4 Tube per pass: 86

Inside diameter: Passes: 1

Utilities: Cooling water Control: Tubular Exchangers Manufacturers Association (TEMA) Standards Tolerances: Heuristic Comment and Drawing: Included DISTILLATION COLUMN Identification: Item: Item Number: T-102 No. Required: 1

Date: 1-Mac-10 By: Wan M. Irman

Function: To recover n-butane and 1-butene

296 Operation: Continuous Materials handled: Feed Bottoms Quantity (kg/hr) 7399.3551 Component (kg/hr) n-butane 1-butene 2-butene Water Carbon Dioxide Oxygen Nitrogen Temperature, °C

15504.4987

Overhead

Reflux

7117.66732

34323.527

7791.010 404.995 1953.007 6899.899 6577.008 6190.235 29851.17 325.594 461.719 434.126 36.997 356.603 499.66 2409.51 13.954

13.524

65.217

6.202

9.253

44.621

116

115

144 Design Data: No. of trays: 31 Feed stage (from bottom):3 Pressure, atm: Top of the column: 3.5 Bottom of the column: 4.5 Reflux ratio: 4.82 Functional height, m: 31.40 Tray spacing, m: 0.91 Material of construction: Stainless Steel 316 Liquid density, kmol/m : 8.23 Vapor density, kmol/m : 0.11 Maximum allowable vapor velocity (superficial), m/s: 0.64 Maximum vapor flow rate, m /s: 760.06 Recommended inside diameter, m: 2.13 Tray efficincy, %: 60 Utilities: Cooling water, steam Control: Reflux ratio, column pressure, number of trays Tolerances: Heuristic Comment and Drawing: Included

115

297

APPENDIX C

EQUIPMENT COSTING

298 Costing Calculations for Isomerazation Process Reactor R-103

The Chemical Engineering Plant Cost Index (CEPCI) of 2001 = 394 The Chemical Engineering Plant Cost Index (CEPCI) of 2010 = 632.6. Pressure = 3 barg Height = 9.58 m Diameter = 3.19 m

Therefore, a) Volume  =  πD2L/4 = (3.14159)(3.19m)2(9.58m)/4 = 76.5661 m3 From Equation A.1 (refer table A.1) log10Cp (2001) = K1 + K2 log10V + K3 (log10 V)2 = 3.3496 -0.2765 log10 (76.5661) + 0.0025 [log10(76.5661)2] = 2.8192 Cp (2001) = 102.8192 =$659.4775 Cp (2010) = $659.4775 (632.6/394) = $1 058.8464 b) From Equation 7.10 with P= 3 barg and D=3.19m

Fp,vessel = = 1.6969

c) From Table A.3 identification number of stainless steel vertical vessel = 20, From Figure A.18, FM = 3.11 d) From Equation A.3 and Table A.4, FBM = [B1 + B2FpFM] FBM = [2.25 + 1.82FpFM] = [2.25 + 1.82 (1.6969)(3.11)] = 11.8548

299 CBM (2010) = $1 058.8464 (11.8548) = 1 USD = RM 3.6996 CBM = RM46 439.9046

Cost for Solvent Recovery Distillation Column (NMP)

The costs of the tower and trays are calculated separately and then added together to obtain the total cost. All the figures, tables and equations are from Analysis, Synthesis, and Design of Chemical Processes, 3rd Edition. The Chemical Engineering Plant Cost Index (CEPCI) of 2001 = 394 The Chemical Engineering Plant Cost Index (CEPCI) of 2010 = 632.6. For the tower: a) Volume  =  πD2L/4 = (3.14159)(2.8m)2( 13.3m)/4 = 81.8949 m3 From Equation A.1 (refer table A.1) log10Cp (2001) = K1 + K2 log10V + K3 (log10 V)2 = 3.4974 + 0.4485 log10 (81.8949) + 0.1074 [log10(81.8949)2] = 4.7665 Cp (2001) = 10

4.7665

=$58 406.78

Cp (2010) = $58 406.78 (632.6/394) = $93 776.9772 b) From Equation A.3 and Table A.4, FBM = [B1 + B2FpFM] FBM = [2.25 + 1.82FpFM] c) From Equation 7.10 with P= 5 bar and D=2.8m

Fp,vessel = = 2.08053

300 From Table A.3 identification number of stainless steel vertical vessel = 20, From Figure A.18, FM = 3.11 So,

FBM = [2.25 + 1.82 (2.08053)(3.11)] = 14.0262

CBM (2010) = $93 776.9772(14.0262) = $1 315 334.638 For the trays:

a) Tray  (tower)  area  =  πD2/4 = (3.14159)(2.8m)2/4 = 6.1572m2 From Equation A.1 : log10Cp (2001) = K1 + K2 log10A + K3 (log10 A)2 = 2.9949 + 0.44651 log10 (6.1572) + 0.3961 [log10(6.1572)2] = 3.9727 Cp (2001) = 103.9727 =$9 390.7440 Cp (2010) = $9 390.7440 (632.6/394) = $15 077.6260 b) From Table A.5, CBM = CpNFBMfq N= 42 fq = 1.0 (since number of trays > 20, Table A.5) c) From Table A.6, SS sieve tray identification number = 61; from Figure A.19,

FBM =1.83 CBM (2010) = $15 077.6260 (42)(1.83)(1.0) = $1 158 866.334

For the tower plus trays, CBM,towe+trays (2010) = $1 315 334.638 + $1 158 866.334 = $2 474 200.972 1 USD = RM 3.6996 CBM = RM9 153 553.916

301 Costing Calculation of Pump (P-104A/B) Basic operation : 1 hour Total Mass :15 320.9696 kg/hr Density,  ρ  :  25  305.939  kg/m3 η       : 80% The Chemical Engineering Plant Cost Index (CEPCI) of 2001 = 394 The Chemical Engineering Plant Cost Index (CEPCI) of 2010 = 632.6.

Qp =

= 0.00016817 m3/s

=

ΔP = 2 bar – 1 bar = 1 bar = 101325 N/m2 Power = ΔP x Qp x = 101325 N/m x 0.00016817 x

= 0.01363 kW=

0.0183hp CBM = Cp x (B1 + B2 FM Fp) x UF CBM = Bare Modules Cost Cp = Purchase Equipment Cost (B1+ B2FMFp) = FBM = Bare Modules Cost UF = Update Factor UF= UF=

= 1.6056

a) From shaft power 0.01363 kW, From Equation A.1 (refer table A.1) log10Cp (2001) = K1 + K2 log10V + K3 (log10 V)2 = 3.3892 + 0.0536 log10 (0.01363) + 0.1538 [log10(0.01363)2] = 2.7154 Cp (2001) = 10

2.7154

=$519.2781

Cp (2010) = $519.2781 (632.6/394) = $833.7445

302 b) From Equation A.3 and Table A.4, FBM = [B1 + B2FpFM], FBM = [1.89 + 1.35FpFM] , At 2 bar, the pressure factor,from Table A.2, Fp = 1 from Equation A.3 Log10Fp = C1 + C2 log10 P+ C3 (log10P)2 c) From Table A.3 identification number of centrifugal pump = 40, From Figure A.18, FM = 4.2 d) So,

FBM = [1.89 + 1.35 (1.0)(4.2)] = 7.56

CBM (2010) = $833.7445 (7.56)(1.6056) = $10 120.2709 1 USD = RM 3.6996 CBM = RM37 440.9541

Cost for Heat Exchanger (E-106) All the figures, tables and equations are from Analysis, Synthesis, and Design of Chemical Processes, 3rd Edition. Total heat transfer area, A = 28.7 m Operating pressure, P = 2 bar The purchase cost of 1 unit shell and tube heat exchanger with heat transfer area of 28.7 m , from figure A.7.4, Cp(2001) = ($700)(28.7) = $20 090 At 2 bar, the pressure factor, Fp = 1 from Equation A.3 Log10Fp = C1 + C2 log10 P+ C3 (log10P)2 For carbon steel shell and tube side, the material factor, FM = 1.0 Multiply Fp x FM = 1.0

303 The value of B1 and B2 for floating-head heat exchanger from Table A.4 is B1 = 1.63 B2 = 1.66 The Chemical Engineering Plant Cost Index (CEPCI) of 2001 = 394 The Chemical Engineering Plant Cost Index (CEPCI) of 2010 = 632.6 So, the bare module cost, CBM from equation A.4 CBM= Cp [B1 + B2FpFM] CBM (2001) = $20 090 [1.63 + 1.66x1] = $66 096.1 CBM (2009) = $ 66 096.1 (632.6/394) = $106 122.8245 1 USD = RM 3.6996 CBM = RM392 612

Cost for Reboiler (E-106)

All the figures, tables and equations are from Analysis, Synthesis, and Design of Chemical Processes, 3rd Edition. Total heat transfer area, A = 9.13 m Operating pressure, P = 2 bar Heat duty = 365.43892 kW The Chemical Engineering Plant Cost Index (CEPCI) of 2001 = 394 The Chemical Engineering Plant Cost Index (CEPCI) of 2010 = 632.6. The purchase cost of 1 unit kettle reboiler with heat transfer area of 9.13 m , from figure A.7.4, Cp(2001) = ($2500)(9.13) = $22 825 Cp(2010) = $22 825 (632.6/394) = $36 647.4492 At 2 bar, the pressure factor, Fp = 1 from Equation A.3 Log10Fp = C1 + C2 log10 P+ C3 (log10P)2 For carbon steel kittle reboiler, the material factor, FM = 1.0

304 Multiply Fp x FM = 1.0 The value of B1 and B2 for kettle reboiler at Table A.4 is B1 = 1.63 B2 = 1.66 So, the bare module cost, CBM from equation A.4 CBM= Cp [B1 + B2FpFM] CBM (2001) = $36 647.4492 [1.63 + 1.66x1] = $120 570.1079 CBM (2009) = $ 120 570.1079 (632.6/394) = $193 585.4067 1 USD = RM 3.6996 CBM = RM716 188.5706

Estimation of Utilities Cost

Utilities cost includes electricity, steam for heating and cooling water for cooling purpose. There are 8000 working hours in this production plant.

Electricity Price per unit = RM 0.26/kWh (TNB) Total power required = 3,263,363.062 kWh/yr Annual electricity cost = 3,263,363.062 kWh/yr X RM 0.26/kWh = RM 848,474.40/yr Water Price per unit = RM 1.20/m3 Total Water required = 158,402.16 m3/yr Annual water cost = RM 190,082.59 Water cooling Price per unit = RM 0.26/1000 kg Amount of water required = 107634.13kg/hr Water required for one month = 77496573.9 kg/month Water cost = 77496573.9 kg/month X RM 0.26/1000 kg = RM 20149.11

305 Steam Price per unit = RM 14.25/1000 kg Amount of water required = 690,892.74 kg/hr Water cost = 690,892.74 kg/hr X RM 14.25/1000 kg = RM 9845.22/yr

95

YEAR

TOTAL EXPENSES

INCOME

ANNUAL EXPENSES

DEPRECIATION

NET PROFIT

TAX (30%)

PROFIT AFTER TAX

NET CASH INCOME

SUMMATION NET CASH INCOME

1

-7447525.716

-7447525.716

-7447525.716

2

-20687571.43

-20687571.43

-28135097.15

3

-60983362.74

-60983362.74

-89118459.89

4

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

0

66,812,654.68

74,008,330.70

-15110129.19

5

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

0

66,812,654.68

74,008,330.70

58898201.51

6

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

0

66,812,654.68

74,008,330.70

132906532.2

7

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

0

66,812,654.68

74,008,330.70

206914862.9

8

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

0

66,812,654.68

74,008,330.70

280923193.6

9

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

334887727.9

10

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

388852262.2

11

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

442816796.5

12

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

496781330.8

13

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

550745865.1

14

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

604710399.4

15

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

658674933.7

16

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

712639468

17

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

766604002.3

18

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

820568536.6

19

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

874533070.9

20

199,559,207.00

125,550,876.30

7,195,676.02

66,812,654.68

20043796.4

46,768,858.28

53,964,534.30

928497605.2

95 Discounted Cash Cumulative

YEAR

NET PROFIT 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68

DCFRR (10%) -6770477.923 -17097166.47 -45817703.04 50548685.68 45953350.62 41775773.29 37977975.72 34525432.47 22886230.47 20805664.07 18914240.06 17194763.69 15631603.35 14210548.5 12918680.46 11744254.96 10676595.42 9705995.836 8823632.579 8021484.162

CUMMULATIVE -6770477.923 -23867644.4 -69685347.43 -19136661.76 26816688.86 68592462.14 106570437.9 141095870.3 163982100.8 184787764.9 203702004.9 220896768.6 236528372 250738920.5 263657600.9 275401855.9 286078451.3 295784447.2 304608079.7 312629563.9

308 YEAR

NET PROFIT 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

YEAR

66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68

NET PROFIT 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68

DCFRR (20%) CUMMULATIVE -6206271.43 -6206271.43 -14366369.05 -20572640.48 -35291297.88 -55863938.36 35690745.9 -20173192.46 29742288.25 9569095.787 24785240.21 34354336 20654366.84 55008702.84 17211972.37 72220675.2 10458688.28 82679363.48 8715573.567 91394937.05 7262977.972 98657915.02 6052481.644 104710396.7 5043734.703 109754131.4 4203112.253 113957243.6 3502593.544 117459837.2 2918827.953 120378665.1 2432356.628 122811021.7 2026963.856 124837985.6 1689136.547 126527122.2 1407613.789 127934735.9

DCFRR (30%) -5728865.935 -12241166.53 -27757561.56 25912373.76 19932595.2 15332765.54 11794435.03 9072642.331 5088833.749 3914487.499 3011144.23 2316264.792 1781742.148 1370570.883 1054285.295 810988.6882 623837.4524 479874.9634 369134.5872 283949.6825

CUMMULATIVE -5728865.94 -17970032.47 -45727594.03 -19815220.27 117374.9357 15450140.47 27244575.51 36317217.84 41406051.58 45320539.08 48331683.31 50647948.11 52429690.25 53800261.14 54854546.43 55665535.12 56289372.57 56769247.54 57138382.12 57422331.81

309 YEAR

NET PROFIT 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

YEAR

66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68

NET PROFIT

DCFRR(40%) CUMMULATIVE -5319661.226 -5319661.2 -10554883.38 -15874544.58 -22224257.56 -38098802.14 19264975.71 -18833826.43 13760696.94 -5073129.493 9829069.241 4755939.749 7020763.744 11776703.49 5014831.246 16791534.74 2611897.396 19403432.13 1865640.997 21269073.13 1332600.712 22601673.84 951857.6516 23553531.5 679898.3226 24233429.82 485641.659 24719071.48 346886.8993 25065958.38 247776.3566 25313734.73 176983.1119 25490717.85 126416.5085 25617134.35 90297.50606 25707431.86 64498.21862 25771930.08

DCFRR(50%) 1 -4965017.144 2 -9194476.192 3 -18069144.52 4 66,812,654.68 14618929.52 5 66,812,654.68 9745953.014 6 66,812,654.68 6497302.009 7 66,812,654.68 4331534.673 8 66,812,654.68 2887689.782 9 66,812,654.68 1403741.379 10 66,812,654.68 935827.5859 11 66,812,654.68 623885.0573 12 66,812,654.68 415923.3715 13 66,812,654.68 277282.2477 14 66,812,654.68 184854.8318 15 66,812,654.68 123236.5545 16 66,812,654.68 82157.70301 17 66,812,654.68 54771.80201 18 66,812,654.68 36514.53467 19 66,812,654.68 24343.02312 20 66,812,654.68 16228.68208

CUMMULATIVE -4965017.1 -14159493.29 -32228637.81 -17609708.29 -7863755.274 -1366453.265 2965081.408 5852771.19 7256512.569 8192340.155 8816225.212 9232148.583 9509430.831 9694285.663 9817522.217 9899679.92 9954451.722 9990966.257 10015309.28 10031537.96

310 YEAR

NET PROFIT 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68 66,812,654.68

DCFRR(60%) CUMMULATIVE -4654703.572 -4654703.57 -8081082.591 -12735786.16 -14888516.3 -27624302.46 11292775.07 -16331527.39 7057984.419 -9273542.967 4411240.262 -4862302.705 2757025.164 -2105277.541 1723140.727 -382136.8141 785287.3284 403150.5143 490804.5803 893955.0945 306752.8627 1200707.957 191720.5392 1392428.496 119825.337 1512253.833 74890.83561 1587144.669 46806.77226 1633951.441 29254.23266 1663205.674 18283.89541 1681489.569 11427.43463 1692917.004 7142.146646 1700059.151 4463.841654 1704522.992