CBB 4023 PLANT DESIGN II
DESIGN OF MALEIC ANHYDRIDE PRODUCTION PLANT
GROUP 6 AMIRAH RAIHANA BINTI HARIS FADZILAH
11885
MARYAM FARZANAH BINTI MOHD FAUZI
11971
MOHAMMAD ILHAM BIN MAT HUSSIN
12004
MOHAMMAD KHAIRULANAM BIN AZEMAN
12005
NUR SYAFIQAH BINTI ABDUL MANAN
12148
CHEMICAL ENGINEERING DEPARTMENT UNIVERSITI TEKNOLOGI PETRONAS SEPTEMBER 2012
CERTIFICATION OF APPROVAL
CBB 4023 PLANT DESIGN II
DESIGN OF MALEIC ANHYDRIDE PRODUCTION PLANT
GROUP 6 AMIRAH RAIHANA BINTI HARIS FADZILAH
11885
MARYAM FARZANAH BINTI MOHD FAUZI
11971
MOHAMMAD ILHAM BIN MAT HUSSIN
12004
MOHAMMAD KHAIRULANAM BIN AZEMAN
12005
NUR SYAFIQAH BINTI ABDUL MANAN
12148
APPROVED BY:
________________________________________ DR. RISZA BINTI RUSLI (Group Supervisor) DATE: 20TH SEPTEMBER 2012
CHEMICAL ENGINEERING DEPARTMENT UNIVERSITI TEKNOLOGI PETRONAS SEPTEMBER 2012
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EXECUTIVE SUMMARY Maleic anhydride is a versatile chemical intermediate used to make unsaturated polyester resins, lube oil additives, alkyd resins, and variety of other products. Maleic anhydride is frequently shortened to MAN. In order to produce MAN, the process involved is by the oxidation of benzene or other aromatic compounds. In this case, we would use the normal butane (n-butane) as the main feed. Regarding to this process, it will be further explain in the next chapter as well as the process route that has been chosen in this project. The main objective of this project is to develop a Maleic Anhydride production plant. The development of plant should consider all the relevant criteria required in order to make the most optimize production plant. Throughout this project, overall document emphasized on the details of the project background, market survey, site feasibility study, conceptual process design, process control, safety and loss prevention, waste treatment facility as well as economic evaluation. Location chosen for MAN production plant is at Kidurong, Sarawak. This is due to the availability of raw materials, utilities and transportation. In United States, MAN production rate is estimated to be about 250,000tonnes/annum. According to the demand, the quantity of MAN is about 223,000tonnes/annum.Thus, the proposed plant design will be justified based on the economic potential of the process, by comparing the price of MAN and price of raw materials needed. Hence, overall process description of this project will be further explained in the next chapters.
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ACKNOWLEDGEMENT
Alhamdulillah, praised to God for giving us an opportunity to complete this Plant Design Project I and II courses after struggling with all the problems and challenges in completing design project for the past several months. There were about fourteen (14) weeks have been given to us in completing the design project in Plant Design Project II (CBB4023) course under the supervision of our keen supervisor, Dr. RiszaRusli. We as the member of this group would like to pass our highest gratitude to Dr. RiszaRusli for all his guidance and continuous supports throughout the semester. He has been a very supportive supervisor and willing to share his knowledge, in order to ensure that we could learn and understand every single thing in this project. Our gratitude is also extended to PDP 2 coordinator, Dr. Rajashekar and DrMurniMelati for their effort in arranging and planning the course structures so that all will be run smoothly. Last but not least, our appreciation is given to our beloved group mates, course mates and also friends, thanks for all supports and motivations, which helps us a lot to make sure that this project ended successfully. Not to forget to those who directly or indirectly involved in giving us the opportunity to learn and work as a team while designing our first plant project.
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TABLE OF CONTENTS
CERTIFICATION OF APPROVAL .................................................................................................................. 1 EXECUTIVE SUMMARY .................................................................................................................................. 2 ACKNOWLEDGEMENT ................................................................................................................................... 3 TABLE OF CONTENTS ..................................................................................................................................... 4 LIST OF FIGURES .............................................................................................................................................. 8 LIST OF TABLES ................................................................................................................................................ 9 CHAPTER 1: INTRODUCTION...................................................................................................................... 10 1.1 PROJECT BACKGROUND ...................................................................................................................... 10 1.2 PROBLEM STATEMENT ......................................................................................................................... 10 1.3 OBJECTIVES ............................................................................................................................................ 10 1.4 SCOPE OF PROJECT ................................................................................................................................ 11 CHAPTER 2: LITERATURE REVIEW ......................................................................................................... 12 2.1 BACKGROUND OF PRODUCT ............................................................................................................... 12 2.1.1
Product Overview: Maleic anhydride (MAN) ................................................................................. 12
2.1.2
History of MAN Production(Timothy R. Felthouse, 2001) ............................................................. 12
2.2 AVAILABLE & FEASIBLE PROCESS ROUTES TO MAN PRODUCTION ......................................... 14 2.2.1
Benzene partial oxidation to MAN (AP-42, CH 6.14: Maleic Anhydride)...................................... 14
2.2.2
N-butane partial oxidation to MAN ................................................................................................ 15
2.2.3
MAN fromphthalic anhydride recovery process ............................................................................. 15
2.3 SCREENING AND SELECTION OF PROCESS ROUTES ...................................................................... 16 2.4 PHYSICAL AND CHEMICAL PROPERTIES ......................................................................................... 17 2.5 COST DATA .............................................................................................................................................. 17 2.6 SITE FEASIBILITY STUDY..................................................................................................................... 19 2.6.1
Introduction .................................................................................................................................... 19
2.6.2
Selection Criteria ............................................................................................................................ 19
2.6.3
Summary of site Characteristic in Each Location .......................................................................... 20
2.6.4
Site Evaluation ................................................................................................................................ 21
2.7 POTENTIAL HAZARDS .......................................................................................................................... 23 2.7.1
Previous Accident On Similar Plant ............................................................................................... 23
2.7.2
Potential Hazards and Control Measures ...................................................................................... 24
2.7.3
Material Safety Data Sheet (MSDS) & Hazard .............................................................................. 26
CHAPTER 3: CONCEPTUAL PROCESS DESIGN AND SYNTHESIS ..................................................... 30 3.1 LEVEL I: PROCESS OPERATING MODE .............................................................................................. 30
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3.2 LEVEL II: INPUT-OUTPUT STRUCTURE ............................................................................................. 31 3.3 LEVEL III: REACTOR DESIGN AND REACTOR NETWORK SYNTHESIS ....................................... 32 3.3.1
Reactor Conversion Selections ....................................................................................................... 33
3.3.2
Preliminary Reactor Mass Balance ................................................................................................ 34
3.3.3
Reactor Type Selection ................................................................................................................... 36
3.4 LEVEL IV: SEPARATION SYSTEM SYNTHESIS ................................................................................. 38 3.4.1
First Process Route ......................................................................................................................... 39
3.4.2
Second Process Route ..................................................................................................................... 40
3.4.3
Process Route Selection .................................................................................................................. 41
3.5 LEVEL V: HEAT INTEGRATION ........................................................................................................... 41 3.5.1
Introduction to Pinch Analysis........................................................................................................ 41
3.5.2
SPRINT Software ............................................................................................................................ 42
3.5.3
Stream Data Extraction .................................................................................................................. 42
3.5.4
Minimum Temperature Difference .................................................................................................. 43
3.5.5
Maximum Process Heat Recovery .................................................................................................. 43
3.5.6
Heat Exchanger Network ................................................................................................................ 46
3.5.7
Energy-Saving Evaluation .............................................................................................................. 47
CHAPTER 4: INSTRUMENTATION AND CONTROL ............................................................................... 48 4.1 INTRODUCTION ...................................................................................................................................... 48 4.2 DESIGN OF PLANT WIDE CONTROL SYSTEM ................................................................................... 48 4.2.1
Procedures ...................................................................................................................................... 48
4.2.2
Drier Control System ...................................................................................................................... 49
4.2.3
Deisobutanizer Control System....................................................................................................... 49
4.2.4
Reactor Control System .................................................................................................................. 50
4.2.5
Heat Exchanger Control System ..................................................................................................... 52
4.2.6
Absorber Control System ................................................................................................................ 53
4.2.7
Stripper Distillation Column Control System ................................................................................. 54
CHAPTER 5: SAFETY AND LOSS PREVENTION ..................................................................................... 56 5.1 HAZARD AND OPERABILITY STUDIES (HAZOP) ............................................................................. 56 5.1.1
Introduction to HAZOP .................................................................................................................. 56
5.1.2
Study Nodes Selection ..................................................................................................................... 56
5.1.3
HAZOP Analysis ............................................................................................................................. 57
5.2 PLANT LAYOUT ...................................................................................................................................... 65 5.2.1
Site Layout ...................................................................................................................................... 65
5.2.2
Non-Process Area ........................................................................................................................... 65
5.2.3
Process Area ................................................................................................................................... 66
5.2.4
Assembly point ................................................................................................................................ 68
5.2.5
Emergency Exit ............................................................................................................................... 69
5.3 PLANT LAYOUT CONSIDERATION FACTORS .................................................................................. 69
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CHAPTER 6: WASTE TREATMENT ............................................................................................................ 72 6.1 INTRODUCTION ...................................................................................................................................... 72 6.2 WASTE MINIMIZATION ......................................................................................................................... 73 6.3 WASTE AUDIT ......................................................................................................................................... 73 6.4 EFFLUENT DISCHARGE STANDARD AND REQUIREMENTS ......................................................... 74 6.4.1
Purpose of Effluent Standards ........................................................................................................ 74
6.4.2
Liquid waste .................................................................................................................................... 75
6.4.3
Gaseous Waste ................................................................................................................................ 76
6.4.4
Solid waste ...................................................................................................................................... 76
6.5 TREATMENT STRATEGY ...................................................................................................................... 76 6.6 SCREENING PROCESS............................................................................................................................ 80 6.6.1
Aerated Grit Removal ..................................................................................................................... 80
6.6.2
pH Stabilizer ................................................................................................................................... 80
6.6.3
Equalization Tank ........................................................................................................................... 80
6.6.4
Coagulation Tank ........................................................................................................................... 81
6.6.5
Clarifier .......................................................................................................................................... 81
6.6.6
Sludge Dewatering.......................................................................................................................... 81
6.6.7
Mechanical Sludge Thickener ......................................................................................................... 81
6.6.8
Sludge Storage ................................................................................................................................ 81
6.6.9
Disinfection ..................................................................................................................................... 82
6.6.10
Biopond ...................................................................................................................................... 82
6.7 GAS TREATMENT ................................................................................................................................... 82 6.8 SOLID HANDLING TREATMENT.......................................................................................................... 85 6.9 SCHEDULED WASTE .............................................................................................................................. 86 CHAPTER 7: PROJECT ECONOMICS AND COST ESTIMATION ......................................................... 87 7.1 INTRODUCTION ...................................................................................................................................... 87 7.1.1
Capital Investment .......................................................................................................................... 87
7.1.2
Total Equipment Cost (TEC) .......................................................................................................... 88
7.1.3
Fixed Capital Investment ................................................................................................................ 88
7.1.4
Estimation of Total Operating Cost ................................................................................................ 89
7.1.5
Gross Profit .................................................................................................................................... 89
7.2 PROFITABILITY ANALYSIS .................................................................................................................. 90 7.2.1
Start-up Period ............................................................................................................................... 90
7.2.2
Depreciation ................................................................................................................................... 90
7.2.3
Cash Flow Estimation ..................................................................................................................... 91
7.2.4
Net Present Worth ........................................................................................................................... 92
7.2.5
Internal Rate of return .................................................................................................................... 92
7.2.6
Rate of Return (ROR) Estimation ................................................................................................... 92
7.2.7
Net Present Value or Worth (NPV) Estimation............................................................................... 92
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7.2.8
Pay Back Time ................................................................................................................................ 93
7.3 DISCUSSIONS .......................................................................................................................................... 93 CHAPTER 8: CONCLUSION & RECOMMENDATION ............................................................................. 94 8.1 CONCLUSION .......................................................................................................................................... 94 8.2 RECOMMENDATION .............................................................................................................................. 95 REFERENCES ................................................................................................................................................... 96 APPENDICES..................................................................................................................................................... 98
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LIST OF FIGURES FIGURE 1: THE MAIN PROCESS TECHNOLOGY FROM SCIENTIFIC DESIGN ...................... 14 FIGURE 2: REACTIONS FOR PARTIAL OXIDATION OF N-BUTANE TO MAN ...................... 15 FIGURE 3: GLOBAL DEMAND OF MAN ACCORDING TO REGION(MALEIC ANHYDRIDE . 18 FIGURE 4: PRODUCTION CAPACITY OF MAN ACCORDING TO REGION, 2000 .................. 18 FIGURE 5: KIDURONG INDUSTRIAL AREA ........................................................................... 23 FIGURE 6: SCHEMATIC PROCESS OF AMMONIA .................................................................. 31 FIGURE 7: HEURISTIC METHOD: ONION MODEL ................................................................. 38 FIGURE 8: FIRST PROCESS ROUTE ........................................................................................ 40 FIGURE 9: SECOND PROCESS ROUTE .................................................................................... 41 FIGURE 10: COMPOSITE CURVE FOR MAXIMUM HEAT RECOVERY: ................................. 44 FIGURE 11: PROBLEM TABLE ALGORITHM BY SPRINT SOFTWARE .................................. 45 FIGURE 12: HEAT EXCHANGER NETWORK .......................................................................... 46 FIGURE 13: SELECTED STUDY NODES CONNECTED TO REACTOR .................................... 57 FIGURE 14: ISOBUTANE STORAGE TANK ............................................................................. 77 FIGURE 15: BLOCK DIAGRAM OF WASTEWATER TREATMENT PLANT ............................. 78 FIGURE 16: FLOW SHEET OF WASTEWATER TREATMENT PLANT..................................... 79 FIGURE 17: NITROGEN SEPARATOR ..................................................................................... 84 FIGURE 18: WASTE OF STREAM 18 (S18) .............................................................................. 85 FIGURE 19: CASH FLOW DIAGRAM ....................................................................................... 91 FIGURE 20: PAYBACK TIME ................................................................................................... 93
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LIST OF TABLES TABLE 1: WORLD MAN CAPACITY BY REACTOR TYPE (MAN WORLD SURVEY, 1992) ... 13 TABLE 2: CONTRIBUTING FACTORS TO OPERABILITY AND ECONOMY ASPECTS .......... 20 TABLE 3: WEIGHTAGE CRITERIA IN DECIDING SITE LOCATION ...................................... 21 TABLE 4: WEIGHTAGE TABLE ON SITES .............................................................................. 22 TABLE 5: ANALYSIS OF POTENTIAL HAZARDS ................................................................... 25 TABLE 6: HAZARD ANALYSIS ............................................................................................... 26 TABLE 7: GUIDELINES FOR BATCH AND CONTINUOUS PROCESS ..................................... 30 TABLE 8: PRODUCTION OF MALEIC ANHYDRIDE FROM N-BUTANE ................................. 33 TABLE 9: PRELIMINARY REACTOR MASS BALANCE .......................................................... 34 TABLE 10: MASS BALANCE WITH RECYCLE ........................................................................ 35 TABLE 11: ADVANTAGES AND DISADVANTAGES OF PACKED BED REACTOR (SMITH) . 36 TABLE 12: ADVANTAGES AND DISADVANTAGES OF FLUIDIZED BED REACTOR ........... 36 TABLE 13: STREAM DATA EXTRACTED ............................................................................... 42 TABLE 14: OPTIMUM T MIN IN DIFFERENT INDUSTRIES (PINCH ANALYSIS).................... 43 TABLE 15: UTILITIES REQUIRED BEFORE HEAT INTEGRATION ........................................ 44 TABLE 16: UTILITIES REQUIRED AFTER HEAT INTEGRATION .......................................... 45 TABLE 17: COMPARISON OF UTILITY REQUIREMENT BEFORE AND AFTER HI ............... 47 TABLE 18: DRIER CONTROL SYSTEM DETAILS ................................................................... 49 TABLE 19 : DEISOBUTANIZER CONTROL SYSTEM DETAILS .............................................. 50 TABLE 20: REACTOR CONTROL SYSTEM DETAILS ............................................................. 52 TABLE 21: COOLER CONTROL SYSTEM DETAILS ................................................................ 53 TABLE 22: HEAT EXCHANGER CONTROL SYSTEM DETAILS ............................................. 53 TABLE 23: GAS ABSORBER, C-301 CONTROL SYSTEM DETAILS ........................................ 54 TABLE 24: DISTILLATION COLUMN CONTROL SYSTEM ..................................................... 55 TABLE 25: STUDY NODES IN HAZOP ANALYSIS .................................................................. 56 TABLE 26: STUDY NODE #1 .................................................................................................... 58 TABLE 27: STUDY NODE #2 .................................................................................................... 60 TABLE 28: STUDY NODE #3 .................................................................................................... 62 TABLE 29: RECOMMENDED MINIMUM CLEARANCE ........................................................... 70 TABLE 30: WASTE STREAMS PROPERTIES ........................................................................... 74 TABLE 31: ENVIRONMENTAL QUALITY (EXTRACT) ........................................................... 75 TABLE 32: PLANT WASTEWATER AND STANDARD B VALUES OF EQA ............................ 76 TABLE 33: METHOD OF REMOVAL ACCORDING TO WASTE COMPONENT ....................... 76 TABLE 34: SUMMARY OF LIMITATIONS FOR THE GASEOUS TREATMENT STRATEGY ... 82 TABLE 35: DISPOSAL METHODS OF SOLID WASTE ............................................................. 85 TABLE 36: TOTAL EQUIPMENT COST ................................................................................... 88 TABLE 37: CAPITAL INVESTMENT FOR START-UP PERIOD ................................................ 90 TABLE 38: CUMULATIVE CASH FLOW .................................................................................. 92
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CHAPTER 1: INTRODUCTION 1.1 PROJECT BACKGROUND The final year design project team has been assigned to design a Maleic Anhydride production plant using mixed butane as the raw material. For this semester, the team needs to incorporate safety aspects, site selection, conceptual design, material and energy balance, heat integration and preliminary economic evaluation in the early design of the plant. Each team is also assigned a lecturer from the Chemical Department as a supervisor to guide the students in designing the plant.
1.2 PROBLEM STATEMENT Maleic anhydride (MAN) is used in the production of unsaturated polyester resins. These laminating resins have a high structural strength and good dielectric properties. These characteristics makes maleic anhydride eligible for a variety of applications in automobile bodies, molded boats, building panels, , chemical storage tanks, lightweight pipe, machinery housings, furniture, luggage, and bathtubs. Maleic anhydride is also used to produce other chemicals including fumaric acid, agricultural chemicals, alkyd resins, lubricants, copolymers, plastics, and succinic acid. The main feed stocks for MAN industry are benzene and normal butane. However, due to its increasing cost of benzene and chemical hazard of benzene, the industry turned to butane as the main feed stock because of its fewer hazards compared to benzene, low cost and abundant source availability. The team is assigned the task of designing a Maleic Anhydride production plant using a prespecified composition of mixed butane as the raw material. The capacity of the plant must be determined based on the demand and worldwide market for Maleic Anhydride.
1.3 OBJECTIVES The objective of this project is to develop a Maleic Anhydride production plant. The plant must be cost-effective, considers all of the desired criteria in addition to dealing with relevant issues. The team needs to recommend the best possible design, which will ultimately convince panel of juries to verify the new plant. The objectives of this design project include the following:
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To integrate chemical engineering skills and knowledge in a detailed design of a chemical plant.
To apply appropriate design codes in a detailed design work
To present a piping and instrumentation diagram (P&ID) and control strategy packages
To perform detailed economic evaluation of the proposed chemical plant
To generate cost effective process options while maintaining operability, safety and environment friendliness of the design.
1.4 SCOPE OF PROJECT The scopes of work for FYDP II are as follows:
Making the necessary decisions, judgements and assumptions in design problems.
Performing the instrumentation and control study
Performing the process design of the major process units.
Performing the mechanical design of the major process units.
Performing the economic evaluation including capital cost estimation and manufacturing cost estimation.
Considering the environmental and safety issues related to the plant. Material safety data sheet (MSDS) for all the chemicals involved must be part of the safety and environmental discussion.
Utilising the blend of hand calculations, spreadsheets, mathematical computer packages, and process simulators to design a process
Preparing the group and individual reports as per standard format and conducting the oral presentations.
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CHAPTER 2: LITERATURE REVIEW 2.1 BACKGROUND OF PRODUCT 2.1.1
Product Overview: Maleic anhydride (MAN)
Maleic anhydride is a versatile molecule that is used in many applications which requires a number of functionalities and properties (Timothy R. Felthouse, 2001). It is considered an excellent joining and cross linking agent due to its three active sites (one double bond and two carboxyl groups). Besides that, due to its cross linking abilities, it is widely used in the manufacturing of unsaturated polyester resins. Maleic Anhydride is also one of the important intermediate in the fine chemical industry, mainly in the manufacturing of agricultural chemicals and additives for lubricating oil. In addition, it also serves as a component for several copolymers in the polymer sectors (Lohbeck, Haferkorn, Fuhrmann, & Fedtke, 2005). In 1928, Diels and Alder worked on a reaction between Maleic Anhydride and 1,4-butadiene and the work was later awarded the Nobel prize in 1950. The starting of the usage of Maleic Anhydride in the pesticide and pharmaceutical industries was because of the studied reaction. Several examples of the specialty chemicals that can be prepared from Maleic Anhydride includes tartaric acid, diethyl and dimethyl succinate, malic acid, glyoxylic acid, diisobutylhexahydrophthalate (DIBE), methyl tetrahydrophthalic anhydride esters and dodecene succinic anhydride. (Lohbeck, Haferkorn, Fuhrmann, & Fedtke, 2005) 2.1.2
History of MAN Production(Timothy R. Felthouse, 2001)
Maleic anhydride was initially commercialized in the early 1930s through the selective oxidation of benzene. The usage of benzene as the feed for the production of maleic anhydride was dominant until 1980s. Several processes were introduced with the common ones that were from Scientific Design. By then, there were also small amount of maleic acid being produced as by-product in the production of phtalic anhydride that can be oxidised into maleic anhydride. However, the usage of benzene started to change and was replaced by n-butane in 1974 because of its toxic effects and economic aspects. The recognition of benzene as hazardous material and deemed carcinogen substance with the rapid increase in the benzene price opened up the search for alternative process technology specifically in the
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United States. Later, the first commercial production of maleic anhydride from butane was established at Monsato’s J. F. Queeny plant in the year 1974. After 1980s, the United States maleic anhydride industries underwent a conversion of feedstock from benzene to butane. But, during the early years, the conversion to butane as feedstock had its limitation whereby the early butane-based catalyst were not active and selective enough for a better conversion of benzene-based plant without a significant loss of capacity production. However, further enhanced catalyst was developed by Monsanto, Denka, and Halcon which led to the world’s first butane-to-maleic anhydride plant which was started up by Monsanto in 1983. The plant incorporated an energy efficient solvent-based product collection and refining system. It was then the largest maleic anhydride production facility and later it undergone debottlenecking project from a capacity of 59,000 tons per year to 105,000 tons per year in 1999. By mid-1980s, United States 100% of maleic anhydride production were using butane as feedstock due to advances in catalyst technology, increased regulatory pressures, and continuing cost advantages of butane over benzene. Meanwhile, Europe has also converted from benzene-based to butane-based maleic anhydride technology starting from the construction of new butane based facilities by CONDEAHunstman, Pantochim and Lonza. The growth in the industry turned to the butane-to-maleic anhydride route, usually at the expense of benzene-based production. Table below shows the worldwide maleic production capacity broken down into fixed-bed benzene, fixed-bed butane, fluidized-bed butane, andphthalic anhydride (PA) co-product. TABLE 1: WORLD MAN CAPACITY BY REACTOR TYPE (MAN WORLD SURVEY, 1992)
1993 Actual Reactor (Feed)
2000 Actual
Kiloton / year,
%
Kiloton / year,
%
Fixed Bed (butane)
369
43.0
704
51.8
Fixed Bed (benzene)
325
37.9
388
28.5
Fluid Bed (butane)
127
14.8
217
16.0
Fixed Bed (PAcoproduct)
37
4.3
50
3.7
Total
858
100.0
1359
100.0
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It can be seen from the table that both fixed-bed and fluidized-bed butane routeshave increased dramatically with the fixed-bed route adding 336 kiloton/year capacity compared to 90 kiloton/year for the fluid-bed process. Only a few newer benzene-based fixed-bed processes have been built with a difference of 63 kiloton/year since the early1980s and the reason it was built was due to limited resource of butane.
2.2 AVAILABLE & FEASIBLE PROCESS ROUTES TO MAN PRODUCTION Generally, there are 3 routes that can be considered for the commercial production of MAN. The 3 routes are given below together with their chemical reactions: 2.2.1
Benzene partial oxidation to MAN (AP-42, CH 6.14: Maleic Anhydride)
FIGURE 1: THE MAIN PROCESS TECHNOLOGY FROM SCIENTIFIC DESIGN
Vaporized benzene and air are mixed and heated before entering the tubular reactor. Inside the reactor, the benzene/air mixture is reacted in the presence of a catalyst that contains approximately 70 percent vanadium pentoxide (V2O5), with usually 25-30% molybdenum trioxide (MoO3), forming a vapor of MA, water, and carbon dioxide. The vapor, which may also contain oxygen, nitrogen, carbon monoxide, benzene, maleic acid, formaldehyde, formic acid, and other compounds from side reactions, leaves the reactor and is cooled and partially condensed so that about 40 percent of the MA is recovered in a crude liquid state. The effluent is then passed through a separator that directs the liquid to storage and the remaining vapor to the product recovery absorber. The absorber contacts the vapor with water, producing a liquid of about 40% maleic acid. The 40% mixture is converted to MA, usually by azeotropic distillation with xylene. Some processes may use a double-effect vacuum evaporator at this point. The effluent then flows to the xylene stripping column where the xylene is extracted. This MA is then combined in storage with that from the separator. Themolten product is aged to allow color-forming impurities to polymerize. These are then removed in a fractionation column, leaving the finished product.
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2.2.2
N-butane partial oxidation to MAN
FIGURE 2: REACTIONS FOR PARTIAL OXIDATION OF N-BUTANE TO MAN
Several process technologies for butane-to-maleic anhydride was designed including Huntsman process, Scientific Design process, Technobell Limited process, Alusuisse maleic anhydride (ALMA) process, and Du Pont moving bed recycle based process. Butane and compressed air are mixed and fed adiabatic reactor, where butane reacts with oxygen to form maleic anhydride. The reaction is exothermic, therefore either a fluidized bed reactor or a packed bed reactor with heat removal to stay close to isothermal. The reactor effluent is cooled and sent to packed bed absorber, where it is contacted with water to remove the light gases and all of the maleic anhydride reacts to form maleic acid. The vapor effluent, which consists of non-condensable must be sent to an after-burner to remove any carbon monoxide prior to venting to the atmosphere. The liquid effluent is then cooled and flashed at 101 kPa and 120°C. The vapour effluent after flashed is sent to waste treatment where else the liquid effluent, is sent to another reactor where maleic acid is broken down to maleic anhydride and water. The reactor effluent is then sent to distillation column where maleic anhydride and water are separated. The distillate from the distillation column is sent to waste treatment (Production of Maleic Anhydride). 2.2.3
MAN fromphthalic anhydride recovery process (MAN as a byproduct of the
production of phthalic anhydride) C8H10 + 7.5O2 C4H2O3 + 4H20 +4CO2∆H = -2518.5 kJ/mole
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The process technology involved in this process route includes LURGI - BASF Phthalic Anhydride Process and Technobell Limited Process. Hot air and vaporized o-xylene are mixed and sent to a packed bed reactor. Most of the oxylene reacts to form phthalic anhydride, but some complete combustion of o-xylene occurs and some maleic anhydride is formed. The reactor temperature is controlled by a molten salt loop. The reactor effluent enters a complex series of devices known as switch condensers. The feed to the switch condensers may be no higher than 180°C; hence, the reactor effluent must be cooled. The net result of the switch condensers is that all light gases and water leave through the top while small amounts of both anhydrides, and the phthalic anhydride and maleic anhydride leave in the bottom. The bottom stream is then purified to obtain maleic anhydride. (Production of Phthalate Anhydride from O-xylene).
2.3 SCREENING AND SELECTION OF PROCESS ROUTES Only two routes are compared which are the benzene route and the n-butane route since the 0-xylene route is only the recovery of maleic anhydride as byproduct. N-butane route is chosen over the benzene route mainly due to benzene’s higher hazard properties compared to n-butane and the rapid increase in benzene’s price over n-butane. In addition, if viewed from the compound’s molecular properties, n-butane route still is a better route. The reason is because there are four carbon atoms in product maleic anhydride per6 carbon atoms in benzene. Therefore, the atom efficiency for the carbon atom is 4/6 x 100% = 66.7%. In the n-butane route, there are four carbon atoms in n-butane per four carbon atoms in maleic anhydride thus giving 100% atom efficiency. If mass efficiency is considered, then the mass of the product is compared to the mass of the raw materials. The molecular weight of maleic anhydride is 98. For the n-butane route, we need 1 mole of n-butane (molecular weight 58) and 3.5 moles of oxygen (total mass of 3.5 x 32 = 112.) Thus the total mass of raw materials needed is 170. The mass efficiency of the n-butane route is therefore 98/170 or 57.6%. By a similar calculation, we can show that the mass efficiency of the benzene route is only 44.4%. As can be seen, n-butane is favourable to benzene in comparison of both atom and mass economy.
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2.4 PHYSICAL AND CHEMICAL PROPERTIES
Properties Formula
Maleic Anhydride (MAN)
Propane
Iso-butane
N-butane
Isopentane
C2H2(CO)2O
C3H8
C4H10
C4H10
C5H12
44.1
58.12
58.12
72.15
0.6011
0.61972
Chemical structure Molecular weight (g/mol) Density (g/ml) Melting point(◦C) Boiling Point(◦C) Flash point(◦C) Appearance
98.06 1.48 52.8 202 103 White crystal
0.584-42 -188
-138
-138
-159.9
-42.1
-11.7
-0.5
27.8
-104 Colourless gas
-87 Colourless gas
-60 Colourless gas
-56 Colourless liquid
2.5 COST DATA In building a production plant, initial capital investment as well as cost for operation and maintenance needs to be considered. According to Silla (2003), there are three components that contribute to the total production cost, which are direct costs, indirect costs and general costs. Direct costs include feedstock supply, utilities supply and labour cost (Silla, 2003). Indirect costs cover expenses like taxes, insurance and plant overhead costs (Silla, 2003). On the other hand, general costs include administrative costs, marketing costs, etc. (Silla, 2003). For the early stage of the market research, studies on the current market of the main product (MAN) and the different feeds are performed. Research on the capital investment for existing production plant is also presented in this paper. MAN is mainly used as a raw material in the production of polyester resins, which are largely used in boating, automobile and construction industries. Recently, a growing market of MAN is observed in the production of butanediol, which is used in the synthesis of plastics and as a 17
textile additive. Apart from that, it is also been greatly used in the manufacture of alkyd resins, a significant material in paints and coatings. Other applications that make use of MAN include manufacturing of agricultural chemicals, lubricant additives and copolymers.
FIGURE 3: GLOBAL DEMAND OF MAN ACCORDING TO REGION(MALEIC ANHYDRIDE, 2011) Africa 1% North America 23% Asia 35%
North America South & Central America West Europe
South & Central America 3%
East & Central Europe Asia Africa
East & Central Europe 4%
West Europe 34%
FIGURE 4: PRODUCTION CAPACITY OF MAN ACCORDING TO REGION, 2000 (TIMOTHY R. FELTHOUSE, 2001)
In relation of the market demand and production capacity according to region, demand in most regions is proportional to the production capacity (Maleic Anhydride, 2011). For an instance, the highest demand of MAN comes from Asia (specifically China), and Asia also has been the topmost producer of MAN (Maleic Anhydride, 2011). In United States, the price of MAN is reported to be between USD 1,922 – 2,055 per tonne in 2011 (Timothy R. Felthouse, 2001). The production capacity in US is reported to be up to 250,000tonnes/year while their demand is 223,000 tonnes/year (Timothy R. Felthouse, 2001).
18
Globally, among companies that supply MAN are Bartek Ingredients Inc., BASF AG, LANXESS Corporation and Huntsman Corporation (Jose, 2008). While in Asia, Changzhou Yabang Chemical Co Ltd, Danyang Chemical Plant, and Thirumalai Chemicals Ltd are the companies with MAN production (Jose, 2008). One company that produces MAN in Malaysia is TCL Industries (Malaysia) SdnBhd (Jose, 2008). TCL Industries Malaysia is a company under Thirumalai Chemicals Ltd that produces worldwide supply of up to 60,000 tonnes per year MAN (Maleic Anhydride, 2011). Considering the world market outlook, China is found to be the main region for future growth. This is due to the rise in demand of unsaturated polyester resin which requires MAN for the production. Besides than Middle East, other countries such as Saudi Arabia and UAE also are showing growth market and an increase in import requirements (ICIS,2002). In a plant design study by Woril Turner Dudley (2012), to build a plant that produce about USD 500,000 for a kmol product/hour MAN requires a capital investment of USD 18,161,381. This value can provide cost estimation in the future plant design for this project.
2.6 SITE FEASIBILITY STUDY 2.6.1
Introduction
Choosing strategic plant location is one of the most crucial decisions needs to be done. The construction of a chemical plant requires a preliminary feasibility study to be done in order to make certain that the proposed Kidurong Industrial Estate is feasible, economically and environmentally. The location of the plant site takes relatively high precedence and it mainly depends on the availability of feedstock, cost of production, marketing of the products, land availability and also the infrastructure. The right location allows maximum profit with a minimum operating cost and allowance for future expansion. 2.6.2
Selection Criteria
Based on the study done in the selecting strategic plant location, there are several factors that should be taken into consideration when undertaking the process of selecting a suitable site. There are two major factors that contribute to the operability and economic aspects of a site location for a plant, which are the primary factor and specific factor.
19
TABLE 2: CONTRIBUTING FACTORS TO OPERABILITY AND ECONOMY ASPECTS
Primary Factors
Specific Factors
Raw material supply for industry
Availability of low cost labor and services
Reasonable land price
Safety and environmental impacts
Source of utilities, such as electricity, water and etc.
Incentives given by government : Pioneer Status Investment Tax Allowance (ITA)
Climate status
2.6.3
Wind Rainfall Temperature Relative Humidity
Effluent and waste disposal facilities
Transportation facilities Local community consideration
Summary of site Characteristic in Each Location
Five major locations are identified to be considered in the site selection for the construction of Maleic Anhydride production plant. The locations are: i)
Kidurong Industrial Area
ii)
Kota Kinabalu Industrial Park
iii)
PasirGudang Industrial Estate
iv)
PengerangIntergrated Petroleum Complex
v)
Kerteh Petrochemical Complex
The characteristics of each location are listed based on the primary and specific factors which had been justified before. Appendix 3 shows the summary of the site characteristics for each location.
20
2.6.4
Site Evaluation
The evaluation of each location is done based on weightage system. Table 3 below shows the range of weighted marks for each identified criteria. The site location is evaluated based on the guidelines. TABLE 3: WEIGHTAGE CRITERIA IN DECIDING SITE LOCATION
Factors Supply of raw material
7-10 Marks
4-6 Marks
Able to obtain large
Source of raw materials from neighbouring states or countries with the distance not exceeding 80km. Uses a pipeline system as well.
supply locally thus saving on import cost Having long pipeline networks for transportation of raw material Price and Area of Land
Land area exceeding 60 hectares Price of land below RM 20 psm
Land area below 60 hectares Price of land more than RM 20 psm
Local Government
Incentives from the local organization of country development Incentives from special company Complete network and well maintained highways, expressways and roads International Airport facilities access to the main location around the world Location near to international port which import and export activities Reliable railway lines to remote areas not accessible by roads
Incentives from the local organization of country development
Incentives
Transportation
21
Good federal road and highway system Limited railway system access More distant from the ports Airport facilities which may not have international flight facilities- only providing domestic flight
0-3 Marks Unable to obtain raw material from close sources with the distance exceeding 80 km Forced to import from foreign countries Uses a pipeline system as well. Land area below 40 hectares Price of land more than RM 30 psm No incentives from the local organization of country development
Average road system No highways or expressway system in close proximity No railway system Very distant from the ports or harbours Distant form the nearest airport more than 100km away
TABLE 4: WEIGHTAGE TABLE ON SITES
Criteria
Kidurong Industrial Site
Kota Kinabalu Industrial Park
PasirGudang Industrial Estate
PengerangIntergrated Petroleum Complex
Kerteh Integrated Petrochemical Complex
Supply of Raw Material
10
8
9
9
10
Price
8
1
9
10
4
Area of Land
10
7
4
9
4
Local Government Incentives
9
8
8
8
8
Transportation
9
8
6
6
10
Workers Supply
10
7
10
9
98
Utilities, water and electricity
8
8
9
7
9
Type of industrial and its location
10
3
5
10
10
Waste water disposal
10
8
10
9
9
Total
84
58
70
86
73
Percent (%)
84
58
70
86
73
Based on the weight matrix, the selected location for MAN plant construction is Kidurong Industrial site. The location is justified to be the most suitable and strategic compared to the others based on all the criteria. Some of the attractive information of the location is as followed:
Kidurong Industrial site is located only 20 km with approximately 15 minutes from Bintulu Town. It is identified as the main industrial core of the whole Bintulu area which is one of the attractions of multinationals oil and gas companies.
It consist of a few established multinational oil and gas company such as MLNG complex, which is deemed as the world largest single gas manufacturing complex apart from Shell and Murphy. 22
FIGURE 5: KIDURONG INDUSTRIAL AREA
Include related infrastructure such as Bintuludeepwater Port, Bintulu Airport and a few options of power supply.
The available MLNG Complex is targeted as the main supply of raw material for Malefic Anhydride production plant.
The land price is cheap and reasonable compared to the other established locations which are around RM77.42. The available land is also huge and sufficient enough for the construction of MAN plant.
2.7 POTENTIAL HAZARDS 2.7.1
Previous Accident On Similar Plant
The plant was located in Indonesia. The fire incident was started on 20th January 2004. This chemical factory was doing processing which involved ammonia and maleic anhydride. The plant employs some 450 people and makes chemicals used in plastics. The blasts were sparked by a fire in several tanks. The fire caused a leak of maleic anhydride from the top of a tank, which ignited, increasing the intensity of the fire. The fire was also threatening some ammonia tanks. Flames were 50 meter high and smoke could be seen several kilometers away. The blaze destroyed at least 5 nearby homes and more than 50 fire trucks were deployed at the height of the fire but only masked firefighters could approach the blaze. The state electricity company cut power to the area. The company shut down pipeline that supplied gas to the industrial complex. The fire was caused by an overheated machine. Fifty-six were injured. Most of them suffered serious burns. The death toll in the devastating fire at a chemical plant increased to three after another victim died from severe burns in 23
hospital. A 36-year-old man sustained burns to more than 90% of his body and at last died after several days. The man was one of the 13 victims being treated in a hospital, who suffered more severe burns in the fire. Three other victims died hours after the fire engulfed the chemical plant. Among them were the plant and production director. Police has evacuated hundreds residents within 1 kilometer of the plant. More than 100 firemen and police battled to put out the blaze. There was the sound of at least five explosions before a column of black smoke rose into the sky. The resulting fire lasted until 2330 hours while thick black smoke covered the sky until 1800 hours. The people have complained of eye irritations from a huge pall of smoke. Hundreds of workers are taking indefinite vacations. It has been discovered that the communities living in the surrounding area, have complained about water contamination in their wells. The water has a foul smell and when used for bathing the community reports it causes itchiness. 2.7.2 Potential Hazards and Control Measures Production of MAN from normal butane requires a process with high pressure and temperature. In addition to that, the reactant used in this petrochemical industry is highly flammable and reactive with oxidizing agent. These conditions made the process plant become relatively hazardous compared to other industries like food and manufacturing industries. Potential hazards are grouped according to its common source, and the hazards effects and control measures are summarized in the table below.
24
TABLE 5: ANALYSIS OF POTENTIAL HAZARDS
Source
Potential Hazards
Chemical Tank, Storage & Transportation
Dangerous chemical reaction, overpressure, temperature above flash point, corrosion of storage tank, release of toxic fume, chemical spillage
Effects
Control Measures
Reactor
Overpressure, overheat, sudden temperature rise
Fire Explosion Air pollution Land pollution
Explosion
Distillation Column
Heat Transfer
Utilities (Electricity, gas, water, etc.)
High pressure
Overpressure, fouling, overflow of cooling or heating medium
Short circuit, ignition
Flooding
Tube rupture Leaking Contamination of liquid Thermal shock
Fire Explosion
Create less conducive, accidentprone workplace
Layout
Location
Environment
Unsafe layout (poor arrangement, limited space to perform safe work)
Natural disaster, approval from local community Untreated waste water
Replace hazardous chemical with a less hazardous alternative Isolation of incompatible chemicals Installation of pressure safety valve Filtration of toxic fume Warning signage/ label for chemicals Store chemical in well-ventilated area Utilize overpressure relief protection Implement control system to prevent overheat and overpressure condition Installation of pressure safety valve Sequence the distillation process for a minimum flow of non-key components Utilize overpressure relief protection Utilize overpressure relief protection Substitution to a less hazardous cooling or heating medium Use control valve to prevent thermal shock Ensure proper connections and maintenance for electrical components Keep flammable material away from ignition source Ensure safe and practical design of process plant, considering safe operational sequence and future expansion Separate process and non-process area. Ensure proper ventilation
High risk of disaster Pollution affecting community (noise, air, etc.)
Selection of site with minimal risk of natural disaster Keep safe distance between the plant site and the local community
Water pollution
Develop a treatment system for waste water
2.7.3
Material Safety Data Sheet (MSDS) & Hazard TABLE 6: HAZARD ANALYSIS
Chemicals
Maleic Anhydride
n-butane
Propane
Flash point
102oC
60.15oC
-104oC
Auto-Ignition Temperature
Fire & Explosion
Reactivity
Symptoms / Effects
477oC
Combustible when exposed to heat or flame. Material in powder form, capable of creating a dust explosion. When heated to decomposition it emits acrid smoke and irritating fumes.
Reactive with oxidizing agents, reducing agents, acids, moisture. Slightly reactive to reactive with metals, alkalis.
Exposure will cause asthma, dermatitis and pulmonary oedema; effects may be delayed. Tumorigen.
286.85oC
450oC
Extremely flammable in the presence of following materials or conditions: open flames, sparks and The product is stable. static discharge and oxidizing materials.
Contact with rapidly expanding gas may cause burns and or frostbite. Acts as a simple asphyxiant.
Explosive air-vapor mixtures may form if allowed to leak to The product is stable. atmosphere.
Higher concentrations may cause dizziness and unconsciousness due to asphyxiant. Liquid can cause burns and frostbite if in direct contact with skin.
Chemicals
iso-butane
isobutene
1-butene
Flash point
-82.8oC
-76oC
-80oC
Auto-Ignition Temperature
Fire & Explosion
Reactivity
Symptoms / Effects
460oC
May be mildly irritating to mucous membranes. At high concentrations, may Flammable liquid and gas Stable. Avoid from high cause drowsiness. At very high under pressure. Form temperature and incompatible concentrations, may act as an explosive mixtures with air materials such as oxidizing asphyxiant and cause headache, and oxidizing agents. agents. drowsiness, dizziness, excitation, excess salivation, vomiting, and unconsciousness. Lack of oxygen can kill.
465oC
Extremely flammable.
In high concentrations may cause asphyxiation. Symptoms may include loss of mobility / Can form explosive mixture consciousness. Victim may not be with air. May react violently aware of asphyxiation. In low with oxidants. concentrations may cause narcotic effects. Symptoms may includedizziness, headache, nausea and loss ofco-ordination.
384oC
Avoid exposure to incompatible materials such Forms explosive mixtures as oxidizing agents, halogens, with air and oxidizing agents. and acids. Avoid to elevated temperatures, and pressures or the presence of a catalyst. 27
Asphyxiant. Moderate concentrations may cause headache, drowsiness, dizziness, excitation, excess salivation, vomiting, and unconsciousness. Lack of oxygen can kill.
Chemicals
Neopentane
Isopentane
n-pentane
Oxygen
Flash point
< -7oC
-51oC
-49oC
-52.2oC
Auto-Ignition Temperature
Fire & Explosion
Reactivity
Symptoms / Effects
450oC
Severe fire hazard. Severe explosion hazard. The vapor is heavier than air. Vapors or gases may ignite at distant ignition sources and flash back. Gas/air mixtures are explosive.
Stable at normal temperatures and pressure. Avoid heat, flames, sparks and other sources of ignition. Minimize contact with material. Containers may rupture or explode if exposed to heat.
High concentrations can cause eye and respiratory mucous membrane irritation and mild narcotic symptoms, or loss of consciousness. Long-term exposure can induce mild dermatitis.
420oC
Flammable in presence of open flames and sparks. Slightly flammable to The product is stable. flammable in presence of oxidizing materials.
260oC
Extremely flammable in presence of open flames and sparks, of heat. Flammable in presence of oxidizing materials. Nonflammable in presence of shocks. Slightly explosive in presence of heat, of oxidizing materials. Nonexplosive in presence of shocks.
Can causes eye and skin irritation. Ingestion may cause central Stable at room temperature in nervous system depression, closed containers under characterized by excitement, normal storage and handling followed by headache, dizziness, conditions. drowsiness, and nausea. Inhalation may cause respiratory tract irritation.
N/A
Oxidizing agent can vigorously accelerate combustion. Contact with flammable materials may cause fire or explosion.
Extremely reactive or incompatible with the oxidizing materials, reducing materials and combustible materials.
28
Hazardous in case of eye contact (irritant), of ingestion, of inhalation. Slightly hazardous in case of skin contact (irritant, permeator).
Breathing 80% or more oxygen at atmospheric pressure for more than a few hours may cause nasal stuffiness, cough, sore throat, chest pain, and breathing difficulty.
Chemicals
Flash point
Auto-Ignition Temperature
Water
N/A
N/A
Carbon Dioxide
Carbon Monoxide
None
Not available.
None
o
Fire & Explosion
Reactivity
N/A
N/A
Non-flammable.
Severe fire hazard. Vapor/air mixtures are explosive. Containers
700 C may rupture or explode if exposed to heat.
29
Stable. Certain reactive metals, hydrides, moist cesium monoxide, or lithium acetylene carbide diammino may ignite. Passing carbon dioxide over a mixture of sodium peroxide and aluminum or magnesium may explode. Stable at normal temperatures and pressure. Incompatibilities withoxidizing materials, halogens, metal oxides, metals, combustible materials, lithium
Symptoms / Effects No acute and chronic health effects.
Moderately irritating to the eyes, skin, and respiratory system.
Harmful if inhaled, blood damage, and difficulty breathing. Causing blisters, frostbite, and blurred vision.
CHAPTER 3: CONCEPTUAL PROCESS DESIGN AND SYNTHESIS 3.1 LEVEL I: PROCESS OPERATING MODE For any chemical process, there are two mode of operation that a plant can choose to operate; batch process and continuous process. The choice of best mode of operation follows these guidelines as in the table shown below: TABLE 7: GUIDELINES FOR BATCH AND CONTINUOUS PROCESS
Guidelines Production rate
Availability
Purpose
Batch
Continuous
Production less than 10 x 106 lb
Production more than 10 x 106 lb
per annum
per annum
Product is a seasonal product
Product is a commodity product
Raw material are limited
Raw material are always available
Suitable for research purposes
Lifetime of Product
Suitable for mass production (profit purposes)
Short
Long
Our selection for the mode of operation of our plant is discussed in detail for each guideline point: i. Production rate Our targeted capacity is 50 000 metric ton per annum, which is bigger than 10 million pound per annum (4 535.9237 metric ton per year). ii. Availability Maleic anhydride is a commodity product, which need to be available all year due to its extensive usage in the industry. iii. Purpose Our maleic anhydride plant obtains profit by its mass production of its product in most economical process route as possible. iv. Lifetime of the product 30
The demand for maleic anhydride is projected to increase and for that, the lifetime for the product is long. v. Operational problem Maleic anhydride is usually produced in a gas state before being cooled and stored as liquid product. With that, it has low operational problem. From all the points discussed above, the best choice of mode of operation for our plant is continuous process as the product is produced in large quantity, the product is a commodity product, it has long lifetime and the process has low operational problem.
3.2 LEVEL II: INPUT-OUTPUT STRUCTURE The process of ammonia is shown below. According to Douglas, to fix the input and output structure, a box is drawn around the process. Then the focus is on what are the feed for the process and what are the product and by-product that comes out from the process:
FIGURE 6: SCHEMATIC PROCESS OF AMMONIA
To complete the input and output structure of the plant, some check question should be answered in order to fix the structure above. The check questions are as shown below: 1. Should we purify the feed stream before they enter the process? 2. Should we remove or recycle a reversible by-product? 3. Should we use a gas recycle and purge stream? 4. Should we not bother to recover and recycle some reactants?
31
We will address each check question in detail in order to synthesize our plant’s input and output structure. Q1. Should we purify the feed stream before they enter the process? By composition, our mixed butane feed consists of approximately 32% unwanted components (such as iso-butane, propane and pentane). These chemicals, without removing it first will affect the reactivity and selectivity of reaction in the reactor. For that reason, the feed should be treated first before being fed into the system. Q2. Should we remove or recycle a reversible by-product? The by-products of MAN production are water, carbon dioxide, and carbon monoxide. These are non-reversible product so there is no need for recycle stream for the by-product. Q3. Should we use a gas recycle and purge system? The process is a high pressure process; the pressure in the MAN synthesis loop is 250kPa. Gas stream that remains after the product being recovered contain some amount of nitrogen. Nitrogen is inert and did not take part in the MAN conversion. They tend to accumulate in the process loop as the reaction take place continuously. This could increase the pressure inside the loop. Therefore, these inert should be taken out and be purged. Q4. Should we not bother to recover and recycle some reactant? Process air, a reactant is being used in excess in this process. Since process air is relatively inexpensive compared to mixed butane, we are not bothered to recover the reactant. However, water collected from the moisture separator vessel can be reheated back to steam using the high energy transferred into the furnace, which can be used in the system. This could save some amount of cost since steam is a type of utility in the plant.
3.3 LEVEL III: REACTOR DESIGN AND REACTOR NETWORK SYNTHESIS Reactors play an important role in all plant especially chemical plant, where the chemical reaction is taking part. As shown by onion model, reactor is said to be the heart of the 32
whole plant. In other words, reactor is the most important equipment among all in the chemical plant. As a result, no reactor means no reaction will occur. Without any reaction, there will be no product produced. From the reactions, mass transfer as well as chemical kinetics and others are calculated. 3.3.1
Reactor Conversion Selections
The chemical reactions involved in the production of Maleic anhydride consist of one main reaction and one side reaction. C4H10 + 3.5O2 C4H2O3+ 4H2O [1] Butane oxygen maleic anhydride water C4H10 + 5.5O2 2CO + 2CO2 + 5H2O [2] Butane oxygen carbon monoxide carbon dioxide water TABLE 8: PRODUCTION OF MALEIC ANHYDRIDE FROM N-BUTANE (BLUM & NICHOLAS, 1982)
An ideal reactor will have a high selectivity on the desired product. In evaluating reactor performance, selectivity is more meaningful to consider than reactor yield. (Smith) Thus, the % conversion with the highest selectivity is selected which is a % conversion of 82.2 with a % yield of 57.6 and a % selectivity of 70. The selectivity for carbon dioxide is 20%.(Slinkard & Baylis, 1975) 33
3.3.2
Preliminary Reactor Mass Balance
Preliminary mass balance calculation for reactor is as below.
in
out Reactor
TABLE 9: PRELIMINARY REACTOR MASS BALANCE
Mass flow rate kg/hr
EP1
Component
In
Out
Isobutane
1287.80
1287.80
n-butane
6560.33
1167.74
Oxygen
39730.87
27884.01
Maleic anhydride
0.00
7158.16
Carbon monoxide
0.00
1168.42
Carbon dioxide
0.00
1835.85
Water
3195.32
10021.49
Nitrogen
130843.36 130843.36
= Product-Reactant = (RM 6 566.43/ton x 51958tonne/yr) – [(RM 2858.34/ton x 56693tonne/yr) = RM 179 130 700.3/year
Assuming that the n-butane can be recycled back as feed, the opportunity to optimize the reactor is as below.
34
TABLE 10: MASS BALANCE WITH RECYCLE
Mass flow rate kg/hr
EP1
Component
In
recycle
Out
Isobutane
1287.80
0
1287.80
n-butane
5392.59
1167.74 1167.74
Oxygen
39730.87
0
27884.01
Maleic anhydride
0.00
0
7158.16
Carbon monoxide
0.00
0
1168.42
Carbon dioxide
0.00
0
1835.85
Water
3195.32
0
10021.49
Nitrogen
130843.36 0
130843.36
= Product-Reactant = (RM 6 566.43/ton x 51958tonne/yr) – [(RM 2858.34/ton x 42709tonne/yr) = RM 219 101 726.9/year
The % increase in profit is = (219101726.9-179130700.3)/179130700.3 X 100 = 22,31% The increase in profit is mainly due to the less usage of n-butane as feed. However, the removal of low boiling point components (nitrogen, oxygen, carbon monoxide, isobutene and carbon dioxide) besides n-butane from the recycle stream will cause the loss of most of the n-butane unless the stream is cooled below the boiling point temperature of n35
butane which is typically below cooling water temperature and is not preferred as refrigeration is required which will be costly. 3.3.3
Reactor Type Selection
The choice selection of type reactor is between two types which is the fixed bed reactor (FBR) and the fluidized bed reactor (FBR). The advantages and disadvantages of PBR are summarized below: TABLE 11: ADVANTAGES AND DISADVANTAGES OF PACKED BED REACTOR (SMITH)
Advantage
Disadvantage
High ratio of heat transfer area to
Difficult to control temperature due to
volume
varying heat load in bed
Can be used for multiphase reaction
Use during careful control of residence
excessive giving hot spots
time
Catalyst temperature can be locally
Off line catalyst regeneration
Have mechanical advantage at high pressure
The advantages and disadvantages of FBR are summarized below: TABLE 12: ADVANTAGES AND DISADVANTAGES OF FLUIDIZED BED REACTOR
Advantage
Good
Disadvantage heat
transfer
and
temperature uniformity
Useful
for
regeneration regeneration)
frequent
Attrition of catalyst can cause carryover Back mixing on the kinetics in the reactor,
catalyst
product destruction and by-product reactions
(online
in the space above the fluidized bed
Vulnerability to large-scale catalyst releases from explosion venting
Require a significant amount of space above the catalyst level to allow the solids to separate from the gases. This exposure of the product to high temperatures at relatively long residence times can lead to side reactions and product destruction
36
The temperature of the reactor is around 400-500 0C due to the highly exothermic reaction involved. At this temperature the components are in gas phase. The removal of heat from the reactor is important and is the key to the reactor type selection. The packed bed reactor (PBR) is chosen over the fluidized bed reactor (FBR) due to several reasons. As shown in the Tables 11 and 12 above, both reactors have a relatively good heat transfer rate. Although the FBR have a slight advantage of good control of the temperature uniformity, control of the hot spots and online catalyst regeneration but the usage of the FBR is quite new in the industry and the understanding of the FBR model is new and not complete as a whole. In addition, the disadvantages of the FBR such as back mixing and catalyst venting also has an impact on the destruction of the product and impact on the environment and safety point of view. Nevertheless, there are suggested strategies to overcome the disadvantages like several proposed patents that claim can control back mixing and also cooling the reactor effluent to prevent the catalyst carryover to foul the heat exchangers. However, these strategies are relatively new and not fully established. Meanwhile, the existence of the PBR is since the production of MAN from benzene before the usage of n-butane. Several strategies have been successfully implemented to overcome the problem with the temperature control and hot spots. Among them are using a small diameter and by using a profile of catalyst through the reactor to even out the rate of reaction and achieve better control. Several reactors can be installed to overcome the off line catalyst regeneration. In addition, the exothermic heat of reaction is removed from the salt mixture by the production of steam in an external salt cooler. Efficient utilization of waste heat from a maleic anhydride plant is critical to the economic viability of the plant. The steam can be used to drive an air compressor, generate electricity, or both. There is also the opportunity to have other equipment to operate at a higher operating temperature giving a better performance of the equipment knowing that there is excess waste. This can cut cost by reducing extra equipment for the separation train.
37
3.4 LEVEL IV: SEPARATION SYSTEM SYNTHESIS Having made initial specification for the reactor, attention is turned to separation of the reactor effluent in the process route screening. This is in line with the heuristic approach for the separation train sequence.
FIGURE 7: HEURISTIC METHOD: ONION MODEL
However, considering that the feed is not of pure n-butane, there is a need to carry out separation before the reactor to purify the feed. A distillation column which is the benchmark for the separating equipment is chosen to separate out the n-butane to increase the purity of the feed. The procedure followed in deciding the process route using the heuristic method is as follows:
Decide on type of separator that will be likely used for the required separation
Decide on the sequencing of the separator to achieve the process requirement 38
The golden rule for separation is to separate heterogeneous mixtures as soon as it forms. However the reactor effluent is of homogeneous mixture. The first choice for separating homogeneous mixture is using the distillation column. Nevertheless, separations using distillation column have circumstances not favoring distillation as below: 1. Separation of low molecular weight materials 2. Separation of high molecular weight heat sensitive material. 3. Separation of components with low concentration. 4. Separation of classes of components 3.4.1
First Process Route
The reactor effluent mainly consists of gases like nitrogen, oxygen, carbon monoxide and carbon dioxide. Thus, separation of the gases which is a class of component cannot be done using distillation column. Alternatives of separation include the creation of another phase within the system by changing the temperature or pressure or by addition of a mass separation agent. This can be done by using an absorber. Since changing the operating parameters require the usage of heating or cooling and the component interested to be separated is the gases, the mass separation agent is used to absorb MAN from the effluent. Water is usually used as the standard mass separation agent (MSA). Alternatively, there are also other organic mass separation agents that can be used but for the first process route, the standard water is used. The usage of other mass separation agent will be discussed in the second process route. The water exiting the absorber contains dissolved MAN and also maleic acid which is the reaction of MAN and water. Water is the major component fraction in the stream exiting the absorber. Another phase separation can be done easily by increasing the temperature to remove a large amount of water. The flash drum is used to separate the different phases after the temperature is increased. After the flash drum, the maleic acid needs to be converted back to MAN before the proceeding to further separate the water. For this purpose, a reactor is used to convert back the acid into MAN at a specified temperature. 39
After the maleic acid reactor, a distillation column can be used to separate water and MAN. However, according to UNIFAC thermodynamic package, MAN and water forms azeotrope which makes the mixture hard to separate. To overcome the hard separation, two distillation columns operating at different pressures are needed to separate the water.
FIGURE 8: FIRST PROCESS ROUTE
3.4.2
Second Process Route
The second process route considers the usage of other MSA in the absorber. Among the commonly used mass separation agent for the process are dibutyl phthalate, dibutyl terephthalate, dimethyl phthalate and diisopropyl phthalate. According to Chen (2002), dibutyl phthalate (DBP) is best used for n-butane processes. Thus, it is chosen as the MSA for the second route. The DBP absorbs most of the MAN and small amounts of water. The MSA then needs to be regenerated through the separation of DBP from MAN and water. Usually, a stripper is used for the regenaration of the MSA. However, DBP can be separated from MAN without using a gas stripper. Hence, a stripper distillation column which is a typical distillation column is used to separate the DBP. Since the amount of water together with MAN is relatively small, the purity of the MAN separated is at 99 weight %. So, further
40
purification will need two more distillation column to further separate the azeotrope mixture which will be uneconomical.
FIGURE 9: SECOND PROCESS ROUTE
3.4.3
Process Route Selection
The second process route is selected over the first process route. This is because more equipment is used in the first route due to the production of maleic acid when water reacts with MAN. In addition, the usage of water as the MSA requires two distillation columns to separate the azeotrope mixtures. On the other hand, the usage of dibutyl phthalate as the MSA simplifies the separation train sequence. Although DBP is more expensive than water, but according US Patent 5069687 and US Patent 4071540, the usage of DBP is more energy efficient due to the avoidance of the evaporation of water and effective in absorbing 99.4% of the MAN. Moreover, due the small amount of water, the MAN exiting the stripper distillation column is already at more than 99 weight % so further purification is not needed.
3.5 LEVEL V: HEAT INTEGRATION 3.5.1 Introduction to Pinch Analysis Pinch analysis is a well-established tool that determines the minimum energy requirement and the optimumdesign of heat exchanger network. This analysis enables a plant design to reach the following goals: Maximizing heat recovery of the system. Minimizing heating and cooling utility consumption. Optimizing the selection of utility sources. 41
Optimizing the trade-off between energy costs and capital cost. 3.5.2 SPRINT Software In integrating heat network into the design, SPRINT software is used. This software may be used to develop composite curve, problem table algorithm, and grand composite curve based on the stream data inserted by a user. This software is also capable of detecting infeasibility of a heat exchanger network design if there is any temperature violation. In short, SPRINT software facilitates the necessary calculations for heat integration. 3.5.3 Stream Data Extraction The first step in performing heat integration is to extract information of streams that require heat duty. These are streams that require change in temperature. The streams that are selected for heat integration will exclude streams for equipment (such as reboiler and condenser). Stream data extracted for this project can be tabulated as follows: TABLE 13: STREAM DATA EXTRACTED
Type
Stream
Ts
Tt
H
CP
(°C)
(°C)
(kW)
(kW/°C)
H1
Hot
Reactor exit
500.0000
125.0000
20,619.0904
54.9842
H2
Hot
Solvent feed
250.0000
35.0000
7,027.2561
32.6849
C1
Cold
Feed vaporizer
25.0000
85.0000
854.8509
14.2475
C2
Cold
Deisobutanizer bottom
75.8398
120.0000
320.0433
7.2473
C3
Cold
Mixer output
116.7270
310.0000
10,187.2383
52.7091
Information on supply temperature (Ts), target temperature (Tt) and heat duty (H) are generally extracted from process flow diagram (PFD) made by iCon simulation. CP is assumed constant at any temperature and is calculated using the following equation:
The streams extracted are categorized into hot and cold streams. A stream is defined as hot when it is surplus in heat (requires cooling) and cold stream when it is deficit in heat (requires heating). 42
3.5.4 Minimum Temperature Difference Temperature difference between hot stream and cold stream is the driving force for heat transfer between the two profiles. Minimum temperature difference, Tmin, is the lowest potential driving force, below which heat transfer is unlikely. Tmin determines the amount of heat recovery in a system. When Tmin is lower, the potential heat recovery from process will be higher. However, according to application experience by KBC Energy Service, the typical Tmin value varies according to industry (Pinch Analysis). TABLE 14: OPTIMUM T MIN IN DIFFERENT INDUSTRIES (PINCH ANALYSIS)
Industrial sector
Typical
Remarks
ΔTmin (°C)
Relatively low heat transfer coefficients, parallel 20 – 40
Oil refining
composite curves in many applications, fouling of heat exchangers
Petrochemical
10 – 20
Chemical
10 – 20
Low temperature process
3-5
Reboiling and condensing duties provide better heat transfer coefficients, low fouling As for petrochemicals Power requirement for refrigeration system is very expensive. ΔTmin decreases with low refrigeration T
Maleic Anhydride production plant is a petrochemical industry. Hence, the optimum Tmin according to the table above is in range of 10°C – 20°C. For this project, Tmin of 10°C is selected so as to maximise heat recovery from the system. 3.5.5 Maximum Process Heat Recovery Initially, the plant is designed so as all the heating and cooling requirements are satisfied by using utilities. The table below indicates the total hot and cold utility requirement before heat integration:
43
TABLE 15: UTILITIES REQUIRED BEFORE HEAT INTEGRATION
Type
Stream
H (kW)
H1
Hot
Reactor exit
20,619.0904
H2
Hot
Solvent feed
7,027.2561
Total cold utility required
27,646.3465
C1
Cold
Feed vaporizer
854.8509
C2
Cold
Deisobutanizer bottom
320.0433
C3
Cold
Mixer output
10,187.2383
Total hot utility required
11,362.1325
By applying pinch analysis, we target on maximizing energy recovery from the process so that the utility requirement can be minimized. The maximization of process heat recovery can be visualized using a composite curve as shown below:
FIGURE 10: COMPOSITE CURVE FOR MAXIMUM HEAT RECOVERY:
With the input of extracted stream data, SPRINT software calculates the maximum heat recovery from process streams and the minimum utility requirement. In composite curve, the maximum process heat recovery is represented by the range of heat load where hot 44
profile and cold profile overlap. The maximum process heat recovery, minimum cold utility (Qcmin), and minimum hot utility (Qhmin) requirements are as follows:
TABLE 16: UTILITIES REQUIRED AFTER HEAT INTEGRATION
H
Utility
(kW)
Minimum cold utility (Qcmin)
16,284.2140
Minimum hot utility (Qhmin)
0.0000
The values of pinch temperature, Qcmin, and Qhmin can also be obtained from problem table algorithm (PTA). Figure below shows the PTA acquired using SPRINT software.
FIGURE 11: PROBLEM TABLE ALGORITHM BY SPRINT SOFTWARE
From the PTA, the pinch temperature,
.
Therefore,
45
3.5.6 Heat Exchanger Network Based on information extracted on PTA, pinch point is located at the highest shifted temperature (495C). This means that the overall system existed below the pinch. Therefore, only analysis of heat integration below the pinch is relevant. Prior to decision on the best heat exchanger network, estimation of the minimum number of heat exchanger units, Nunit, can be made using the following equation:
Where From the equation, During pairing of the streams, two rules must be obeyed to ensure the network is feasible. The rules are:
The T between a pair of hot stream and cold stream must always be ≥ 10C.
CP rule (i.e. CPhot≤ CPcold for above pinch and CPhot≥ CPcold for below pinch) must not be violated unless the pair is away from pinch.
Figure below illustrate the heat exchanger network designed for this plant. Tpinch hot, 500°C 10,432 kW
H1 H2
314.72C
500
125 240.21 C
250
214.05C
35
5,852 kW
C1
85
C2
120
25
855 kW 75.8 320 kW
C3
310
116.7 10,187 kW
Tpinch cold, 490°C FIGURE 12: HEAT EXCHANGER NETWORK
After creating heat exchanger network, it can be summarized that the total number of heat exchangers for the plant is 5 units. Process heat recovery can be sum up to 46
11,362.1325kW. The amount of cold utility requirement is 16,284.2140 kW, while no hot utility is required. 3.5.7 Energy-Saving Evaluation Comparison of heat utility requirements before and after heat integration (HI) can be summarized into the following table: TABLE 17: COMPARISON OF UTILITY REQUIREMENT BEFORE AND AFTER HI
Utility Requirement
Heat Duty (kW) % Saving Before HI
After HI
Hot utility, Qh
11,362.1325
0.0000
100.0%
Cold utility, Qc
27,646.3465
16,284.2140
41.1%
Total
39,008.4791
16,284.2140
58.3%
Percentage saving can be calculated using the following formula:
Referring to the table above, it is observed that there is a huge reduction in utility requirement after pinch analysis is applied. The plant can eliminate 100% of the requirement for hot utility, while the amount of cold utility can be reduced up to 41.1%. As an overall, heat integration implemented can minimize the plant’s operating cost by minimizing utility requirement, as well as reducing capital cost by minimizing the number of heat exchangers during heat exchanger network design.
47
CHAPTER 4: INSTRUMENTATION AND CONTROL
4.1 INTRODUCTION A safe mode operating chemical plant requires a good control system around the equipment by installation of relevant instrumentations in achieving its target production. This is to ensure that the operation of the plant could be conducted in the most economical way and avoid accidents that may lead to the upset value of the production. The objectives of designing the control system strategy are:
To have a safer operation plants and avoid accidents such as explosion
To maintain the operational condition of the unit operation at their own respective condition
To control the process that is in line with the production rate
To maintain the product purity as high as possible
To avoid excess usage of heating and cooling utilities
4.2 DESIGN OF PLANT WIDE CONTROL SYSTEM 4.2.1 Procedures The four major steps in designing the plant wide control system are; i.
The overall specification for the plant and its control system are stated.
ii.
The control system structure, which includes selecting controlled, measured and manipulated variables, product quality, handling operating constraints, is developed.
iii.
Detailed specification of all instrumentation, cost estimation, evaluation of alternatives and the ordering and installation of unit operation are available.
iv.
Design and construction of the plant and plant tests (startup, operation at design conditions and shutdowns) are known.
However, for this plant, the steps taken are only; i.
Identification of controlled, manipulated and disturbance variables are developed. 48
ii.
Appropriate control strategies are implemented.
The Process &Instrumentation Diagram (P&ID) for the whole plant is attached in the Appendix 2.
4.2.2
Drier Control System
The purpose of the control system is to maintain the desired water content in the air which is fed into the reactor. The control system consists of pressure and level controller. The drier is a flash vessel whereby flashing of the air occur at a specified temperature and pressure to remove water. TABLE18: DRIER CONTROL SYSTEM DETAILS
Controlled variable
Manipulated variable
Type of controller
Set Point
Pressure of the flash
Vapour flow rate
Feedback Control
Pressure :
vessel
1480kpa
Level of the flash
Air inlet flow rate
Feedback Controller
Level: 70%
drum
4.2.3
Deisobutanizer Control System
The control system in the deisobutanizer column is important in order to have an accurate separation between the components and to ensure good ratio of purity. Basically the process is applying the distillation concept; the only different is the different take off point at the trays. The control system must be able to cause the average sum of the product streams to be exactly equal to the average desired stream. It must also be able to maintain the desired concentration of products at the bottom and the top streams. Control requirements for deisobutanizer, K-1 are:
Pressure in the column Feedback control is used to control the operating pressure at the top of the column to avoid overpressure.
Level of the liquid at the bottom of the column
49
The liquid level is controlled by manipulating the outlet of the bottom product flow rate using feedback control system.
Temperature of the column The reflux temperature is controlled by manipulating the inlet flow rate of low pressure steam (LPS) entering the reboiler by using the feedback control.
Reflux temperature The temperature difference must not exceed the critical temperature difference. The temperature indicator and controller installed at the condenser outlet controls the temperature by manipulating the cool water supply (CWS) line.
Reflux ratio A flow indicator and controller is installed at the reflux line cascaded with flow indicator and controller of the distillate line to control the flow of the reflux back into the column. Ratio control will maintain the ratio of the reflux line and distillate line at the set point. TABLE19 : DEISOBUTANIZER CONTROL SYSTEM DETAILS
Controlled variable
Manipulated variable
Type of controller
Set Point
Pressure in the column
Top outlet vapor flow
Feedback control
Pressure:
rate Bottom liquid level
1100 kPa
Bottom liquid outlet
Feedback control
Level: 80 %
Feedback control
Temperature:
flow rate Bottom temperature
Steam inlet flow rate
75.8 oC Reflux temperature
Cooling water supply
Feedback control
55.4 0C
flowrate Reflux ratio 4.2.4
Temperature:
Outlet gas flow rate
Ratio control
Reflux ratio: 1.8
Reactor Control System
For the control system design at the reactor, C-1, several requirements have to be made;
Feed flow rate of the two inlet streams into the reactor The flow rate into the reactor has a certain ratio value for the ideal reaction to occur or otherwise, accidents such as explosion may occur at any time during the 50
operation. The ratio is set at least 1.7% oxygen. Therefore, these two inlet streams have to be adjusted at this desired value so that accident will not occur by controlling it with a ratio controller. Both are feedforward controller and are interconnected with a ratio controller to ensure that the desired ratio of the feed stream can be obtained.
Temperature of the reactor On top of that, the temperature inside the reactor is controlled by manipulating the bypass supply stream of the refrigerant and the bypass feed flow rate to the reactor using feedback controller. The flow rate of molten salt (MS) will determine and change the temperature inside the reactor until it reaches the desired set point. In addition, for further safety reason, the feed flow rate will be decreased if the temperature goes too high. Additional safety features includes the installation of an interlock system and emergency shutdown system if the temperature gets extremely high. The system is activated at the temperature of 6500C and when it is activated, the block valve will close the flow of feed; the feed is bypassed and the blow down valve will be opened to purge the content in the reactor to prevent contain the high temperature.
Pressure of the reactor The pressure of the reactor also has to be controlled just as the same priority as the temperature in the reactor. The feed are of gaseous phase and are easily influenced by the adjustment of the temperature as well as the pressure with assumption that the gas is an ideal gas. Furthermore, these two parameters are essential in the production of the maleic anhydride as the rate of the reaction is directly dependent on the temperature and pressure of the reactor. The pressure is controlled by regulating the outlet gas at the top of the reactor. For safety reasons, pressure relief valve, PSV is installed should the pressure increased tremendously in the reactor. Accidents may occur if the pressure exceeds the desired value. The catalyst used is also being control in sense of temperature by measuring and manipulating the pressure so that the ideal condition could be achieved.
51
TABLE20: REACTOR CONTROL SYSTEM DETAILS
Controlled Variables
Manipulated Variables
Ratio of inlet flow rate
Air flow rate
Type of Controller Ratio Control
Set Point Flow rate ratio = S8/S6 = 0.017
Temperature of the
Bypass MS and bypass
High selector with
reactor
feed
split range and
Temperature: 500 0C
cascade control Temperature: 6500C
Total bypass feed and
None (use interlock
open blow down valve
and tripping system)
Temperature of the
Start-up heater flow
Feedback control
Temperature: 4000C
reactor (for start-up
bypass
Feedback Control
Pressure: 250 kPa
reactor) Pressure of the reactor
4.2.5
Reactor outlet flow rate
Heat Exchanger Control System
The purpose of performing a control system for heat exchanger is to maintain the desired temperature of its outlet stream either by providing heating or cooling utility. There are two types of heat exchanger control system, depending on whether the heat exchanger has been integrated or not.
Cooler (Single stream heat exchanger) There is only one process stream which is experiencing heating or cooling and the temperature of the stream is controlled by manipulating the flow rate of utilities stream.
Double stream heat exchanger The heat exchanger has been integrated. Therefore, there were two process streams connected to the heat exchanger. Both process streams will experience either heating or cooling process. Bypass stream is connected from inlet stream to the outlet stream of the cold stream. Temperature indicator and controller are 52
installed at the outlet of the hot stream and are cascaded with the flow indicator and controller which is installed at the bypass stream. The temperature of the cold stream is controlled by manipulating the flow rate of bypass stream. TABLE 21: COOLER CONTROL SYSTEM DETAILS
Cooler
W-6
W-9
Controlled
Manipulated
Type of
Variable
Variable
Controller
Reactor effluent
Cooling water
temperature
supply flow rate
Solvent recycle
Cooling water
temperature
supply flow rate
Feedback Control
Set point Temperature: 125 0C
Feedback Control
Temperature: 35 0C
TABLE 22: HEAT EXCHANGER CONTROL SYSTEM DETAILS
Heat
Controlled
Manipulated
Type of
Exchanger
Variable
Variable
Controller
W-4
Outlet temperature Bypass flow rate
Cascade Control
Outlet temperature Bypass flow rate
Cascade Control
Temperature: 35 0C
of W-4
4.2.6
Temperature: 204.2 0C
of pump W-1
Set point
Absorber Control System
Absorber is used to absorb the maleic anhydride from other gases. The requirements for absorber control system are as below:
Level of liquid in the column The level of liquid at the bottom of the absorber is maintain at desired level by controlling the outlet flow of rich solvent using feedback controller.
Pressure at top of the column Feedback control is used to control the operating pressure at the top of the column to avoid overpressure. The pressure at the top of the absorber is maintained at desired condition by controlling the outlet flow of off-gases. For
53
safety reasons, pressure relief valve, PSV is installed should the pressure increased tremendously in the column.
Feed flow rate of the two inlet streams into the absorber The purpose the control system is to ensure the feed streams is at desired flow rate. The feed to absorber are outlet of reactor cooler, W-3 and the recycled solvent dibutyl phthalate from the outlet of heat exchanger, W-4. The ratio for both streams must be controlled to ensure the absorption for maleic anhydride is effective. Ratio control is implemented to ensure that sufficient amount of dibutyl phthalate is fed into the absorber depending on the flow rate ratio set for both inlet streams. TABLE 23: GAS ABSORBER, C-301 CONTROL SYSTEM DETAILS
Control Variable Liquid level at bottom of absorber
Manipulated Variable Bottom outlet flow rate
Type of Controller Feedback Control
Pressure in the Absorber Ratio of inlet flow Rate
Off-gas outlet flow rate
Feedback Control
4.2.7
Ratio Control
Flow of dibutyl phthalate
Set Point Level: 80 % Pressure: 200kPa Flow rate ratio: S29/S16 = 5.71
Stripper Distillation Column Control System
The control system in the distillation column is important in order to have a sharp separation between the components in the incoming feed. The control system of distillation column is controlled based on three purposes: Material balance control - The column control system causing 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 needed to ensure a smooth and safer transition from the old production level to the newly desired production level. Its purpose is to direct the control action in 54
such a way as to make the inflows equal to the outflows and achieve a new steady-state material balance for the plant. Product quality control - To maintain the desired concentration of the products at the bottom and the top of the column. Satisfaction of constraints. For safety purposes, satisfactory operation of the column, certain constraints must be understood and followed, for example: - The column shall not flood. - Column pressure drop should be low enough to maintain the efficiency of the column operation in order 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 so that overload of reboiler or condenser heat-transfer capacity can be avoided. - Column pressure should not exceed a maximum permissible limit. TABLE 24: DISTILLATION COLUMN CONTROL SYSTEM
Controlled Variable Pressure in the column Bottom liquid level Bottom temperature Reflux drum level Reflux temperature Reflux ratio
Manipulated Variable Top outlet vapor flow rate Bottom liquid flow rate Steam inlet flow rate Reflux solution flow rate Cooling water supply flow rate Reflux flow rate & MA product flow rate
55
Type of Controller Feedback Control Feedback Control Feedback Control Feedback Control Feedback Control Ratio Control
Set point Pressure: 120 kPa Level: 80 % Temperature: 250 0C Level: 50 % Temperature: 85 0C Ratio: 1.8
CHAPTER 5: SAFETY AND LOSS PREVENTION 5.1 HAZARD AND OPERABILITY STUDIES (HAZOP) 5.1.1 Introduction to HAZOP HAZOP, according to Dunn (2009), is a structured procedure intended to proactively identify equipment modifications and/or safety devices required in order to avoid any significant danger as a result of equipment failure.HAZOP study focuses on each pipeline and vessel shown on the respective flow sheet or line diagram. Any possible deviations from normal operating conditions are studied.As a result, HAZOP study enables investigation on potential mal-operation, as well as identifies their root causes, consequences and corrective actions. 5.1.2
Study Nodes Selection
For this HAZOP studies, the selection of study nodes is based on nodes that contain highly hazardous materials and critical process conditions. With reference to preliminary hazard analysis conducted in the earlier stage of plant design, it is found that reactor is one of the high-risk equipment. Due to its high temperature application and variation in reaction activity, reactor is classified as equipment with critical condition. Furthermore, the stream contains mixture of gaseous n-butane and air at high temperature, which may lead to explosion if overpressure. Therefore, three study nodes connected to reactor (C-1) is selected to be studied in this HAZOP analysis. The study nodes are shown in the table and figure below. TABLE 25: STUDY NODES IN HAZOP ANALYSIS
No.
Study node
1
Inlet to reactor (C-1) – Stream 11
2
Outlet to reactor (C-1) – Stream 12
3
Inlet of molten salt loop to reactor
56
3
Boiler (For steam generation)
Molten salt inlet 1
Molten salt outlet Reactor
2
C-1 Stream 12 (Product: MAN, water, CO, CO2)
Stream 11 (Reactant: nbutane & air)
FIGURE 13: SELECTED STUDY NODES CONNECTED TO REACTOR
Further understanding of the process flow can be done by referring to Process Flow Diagram (PFD) or Process Instrumentation and Control Diagram (P&ID) as attached in Appendices. 5.1.3 HAZOP Analysis The following tables (Table 26 - Table 28) illustrate findings from HAZOP analysis.
57
TABLE 26: STUDY NODE #1
HAZOP STUDY RECORD SHEET
PROJECT : MALEIC ANHYDRIDE PLANT
Date : 20 October 2012
MAJOR UNIT : REACTOR System : MAN Production Item
Guide word
Deviation
Study node: Inlet to reactor (C-1) – Stream 11 Possible consequences
Possible causes
Safeguard/ Action required
1A1
Flow
More
Upsetting process
Excess air
Install high flow alarm
Uncontrolled reaction
Excess n-butane
Install dry air control valve after
Excessive heat
Valve fully open
compressor Operator to continuously monitor pipe/valve condition and manually change valve position
1A2
No
No reaction
Pipe rupture
Install feed flow meter
No product
Valve fully closed
To be included in biweekly routine
Blockage 1A3
Less
checkup procedure by Inspection Team
Less pressure
Pipe leakage
Install low flow alarm
Less favorable product.
Valve partially closed
Operator to continuously monitor pipe/valve condition and manually change valve position
58
TABLE 26 (CONTINUED)
Item
Guide word
Deviation
Possible consequences
Possible causes
Safeguard/ Action required
1B1
Pressure
More
High velocity of
Valve failure
High pressure alarm
Overheat
Installpressure safety valve (PSV)
Dryer failure
Install high temperature alarm
Less reaction
Leakage
Insert safety valve
Less favorable product
Compressor failure
Install low pressure alarm
Uncontrolled reaction
Pressure increase
Install temperature sensor
Low pressure
Install temperature sensor
Presence of oxidizing
Insulation of pipe and equipment
reactant Overpressure High temperature 1B2
1C1
Less
Temperature
More
Mechanical failure May cause explosion 1C2
Less
Less pressure, less favorable product.
1D1
Ignition
-
Fire, sparks
substance Temperature rise
59
Temperature sensor
TABLE 27: STUDY NODE #2
HAZOP STUDY RECORD SHEET
PROJECT : MALEIC ANHYDRIDE PLANT
Date : 20 October 2012
MAJOR UNIT : REACTOR System : MAN Production Item
Guide word
Deviation
Study node: Outlet to reactor (C-1) – Stream 12 Possible consequences
Possible causes
Safeguard/ Action required
2A1
Flow
More
Upsetting process
Increase in temperature
Install high flow alarm
Increase demand for
Uncontrolled reaction
Regular monitoring by Control Team
Pipe rupture, valve
To be included in biweekly routine
solvent 2A2
No
No product
fully closed, blockage No reaction
checkup procedure by Inspection Team Regular monitoring by Control Team Operator to continuously monitor pipe/valve condition and manually change valve position
2A3
Less
Less product form
Low reaction rate
Install low flow alarm
Decrease in
Pipe leakage
To be included in biweekly routine
temperature
Valve partially closed
checkup procedure by Inspection Team Regular monitoring by Control Team
60
TABLE 27 (CONTINUED)
Item
Guide word
Deviation
Possible consequences
Possible causes
Safeguard/ Action required
2B1
Pressure
More
Low absorption rate
Uncontrolled reaction
Install pressure safety valve (PSV) Install high pressure alarm
2B2
Less
Less flow rate to the absorber
Low reaction rate
Install low flow alarm
Pipe leakage
Operator to monitorpipe/valve
Agitate absorption
condition Regular monitoring by Control Team
process Take longer time 2C1
Temperature
More
Increase in utility
Uncontrolled reaction
Install temperature sensor Install auto-control valve for utility
2C2
Less
Less pressure
Low reaction rate
Take longer time for absorption process
61
Install temperature sensor
TABLE 28: STUDY NODE #3
HAZOP STUDY RECORD SHEET
PROJECT : MALEIC ANHYDRIDE PLANT
Date : 20 October 2012
MAJOR UNIT : REACTOR System : MAN Production Item
Guide word
Deviation
Study node: Inlet of molten salt loop to reactor Possible consequences
Possible causes
Safeguard/ Action required
3A1
Flow
More
More heat absorbed
Valve fully open
Temperature in reactor
Install high flow alarm Operator to manually change valve
decrease
position
Favor undesired reaction 3A2
No
Rise in reactor temperature
Pipe rupture
Install safety valve
Valve fully closed
Pipe/valve operation to be included in
Explosion
biweekly routine checkup procedure by
May lead to auto-
Inspection Team Operator to continuously monitor
ignition
pipe/valve condition and manually change valve position
62
TABLE 28 (CONTINUED)
Item
Guide word
Deviation
Possible consequences
Possible causes
Safeguard/ Action required
3A3
Less
Rise in reactor temperature
Pipe leakage
Install low flow alarm
Valve partially closed
Install safety valve
Explosion
To be included in biweekly routine
May lead to auto-
checkup procedure by Inspection Team Operator to continuously monitor
ignition
pipe/valve condition and manually change valve position 3B1
Pressure
More
System unit rupture
Utility (boiler) failure
Install pressure safety valve (PSV)
Backflow into solvent
Blockage due to crystal
Install high pressure alarm
supply
formation
Boiler operation to be included in biweekly routine checkup procedure by Inspection Team
3B2
Less
Reduction in flow rate
Large heat loss in
of salt Longer cooling time for
utility (boiler) system Pipe leakage
reactor
Install low flow alarm Boiler/pipe operation to be included in biweekly routine checkup procedure by Inspection Team
63
TABLE 28 (CONTINUED)
Item
Guide word
Deviation
Possible consequences
Possible causes
Safeguard/ Action required
3C1
Temperature
More
Less effective in
Utility (boiler) failure
Install temperature sensor Boiler operation to be included in
cooling down reactor’s temperature
biweekly routine checkup procedure by Inspection Team
3C2
Less
Heat capacity is higher
Large heat loss in
More heat is absorbed
utility (boiler) system
Temperature in reactor
Install temperature sensor Boiler operation to be included in biweekly routine checkup procedure by Inspection Team
decrease Favor undesired reaction
64
5.2 PLANT LAYOUT 5.2.1
Site Layout
This Maleic Anhydride production plant will be located in Kidurong Industrial Area, Sabah. It is planned to occupy area of 350 meter x 300 meter (105,000 m2). The overall site layout is shown in Appendix. Several factors have been considered in laying out the site. The process units and ancillary building should be laid out to give the most economical flow of materials and personnel around the site. In term of safety, process area is located at enough distance from the place where there are a lot of personnel. Basically, the site layout can be divided into two parts:
Non-Process area
Process Area 5.2.2
Non-Process Area
As suggested by Kirk-Othmer (1997), non-process area should be located at a distance of at least, 60 meters from processing area. This is important to avoid any undesired incident due to explosion or fire from the process zone. Among the buildings or units in the non-process area are:
Guard post
Guard posts are located at the entrance of the site. The security checkpoints are important to ensure unauthorized access into the plant. There are 3 guard posts in this site: o Main entrance guard post – to control the flow in and out of personnel or cars between the site and public area. o Process area guard post – security check to ensure that there is no hazardous or undesired materials being brought into the process area. It is a common practice in some places where personnel should have a ‘safety-briefing’ card to be allowed to enter the process site. o Contractor’s entrance guard post – contractors will enter the process area at different entrance with their heavy transport such as crane and lorry.
Administration building 65
Administration office building is sited far away from the process area to avoid any explosion and fire hazard since this is the place where a lot of personnel involve. It is also near to the cafeteria and ‘surau’.
Cafeteria
Cafeteria provides meals for the employees and visitors.
Clinic
The location of clinic has been chosen so that it can be reached easily either from the process area or non-process area. It offers emergency and fast treatment to the injured employees before being sent to the nearest hospital for further treatment. 5.2.3
Process Area
Process area is the heart of this plant. It is a hazardous area since it deals with a lot of chemicals. Arrangement of main process site as well as other ancillary buildings was done carefully. Below are the units and buildings in the Process Area:
Unit 1
This is where the n-butane is being separated from the feed mixture in obtaining a stream of high content n-butane. It consists of a deisobutanizer column, compressor and mixer.
Unit 2
Oxidation process occur in vapor phase. The feed is contacted with the Vanadium Phosphorus Oxide catalyst in a packed bed system. MAN is formed as the main product in the process with the production of carbon monoxide and carbon dioxide as the by product.
Unit 3
MAN purification is a unit to further purify the MAN from other by-products. The main process goes through the absober Column where MAN is absorbed from the incoming stream followed by a stripper unit for DBP regeneration to be recycled.
Utilities
This unit will supply cooling water, low pressure steam, plant and instrument air and some other utilities to the main process unit. Its location is perfectly suitable to give the most economical run of pipe to and from the process unit. 66
Pump house
The pump house contains pumps used to control the material stream flow between the process units with the storage farm.
Wastewater Treatment Unit
The wastewater effluent from the process unit will be sent to Wastewater Treatment Unit to be treated before being released to the environment. It is located adjacent to the main process unit so that the wastewater effluents from the process units can be channeled to it without needing long piping to be transferred.
Storage farm
Storage farm consists of some big tanks. These tanks will store the raw material from supplier before being processed and also stores the product before being exported. Storage farm is located away from the major processing unit to avoid explosion hazard.
Central control building
All the control valves for the whole process area will be controlled and monitored from this central control building. Even it is near to the main process area, it still can be considered as a safe place since it is provided with explosion proof doors and very thick concrete walls. In case of emergency occurs in the plant, control room will be the assembly point in the process area.
Fire station
There are two fire stations provided in the process area. One is located near the main process unit and the other one is placed near the storage farm. These arrangements are made so that faster action could be taken in case of fire-emergency in these two most hazardous areas.
Operator station
Operator station is adjacent to the central control building.
Laboratory
Quality of feed and product should be taken into considerations. Laboratory is the place where the sample (both feed and product) is tested and analyzed to determine whether it meets the specifications or not. All the result will be sent to the control room and some 67
adjustments in controlling will be made, if needed. Thus, the distance between laboratory and control room is not too far. Laboratory staffs will also perform an analysis regarding waste of the process before being channeled to nearby environment.
Chemical storage
This unit stores vessels containing chemical substance, lubricants, and catalyst pellet used for the process. Thus, it is located to the process unit.
Flare Area
Flare is used to burn all excess gas that is emitted from the process units as well as to burn some of the waste gas from waste treatment area. In our plant, the flare stack is located at the middle of 100m x 100m area, to give a radius of 50 meters from other sites and buildings, and meet the statement from Kirk-Othmer (1997), which stated that the minimum safety range from flare to unit operation and storage farm is in radius of 60 meters.
Warehouse
Warehouse stores all the equipment’s spare parts. It is placed near to the workshop to ease the maintenance job.
Expansion Site
There are some free areas allocated for the future plant expansion. They occupy enough space for further expansion, whether for process reaction or producing the plant’s own utility such as cooling water and steam. 5.2.4
Assembly point
For the whole site in this vinyl chloride plant, there are a few zones that have been identified to be as assembly area. In the site layout shown in the appendix, all the assembly areas are represented by small triangles. These are the focal points for every personnel to gather in case of emergency occur, and the assembly areas are located in both process area and nonprocess area. For the non process area, the assembly points are determined to be in front of the car park, administration building and warehouse. For the process area, the assembly points are in the control room, near the process area main entrance as well as beside the fire
68
station. However, control room is the best assembly point since it is built with special safety features with thick concrete wall and explosion-proof door. 5.2.5
Emergency Exit
It is a common practice to have some alternatives way to exit from the chemical plant. In this site, there are three emergency exits available, two of them are provided in the process area while the other one is near to the workshop building.
5.3 PLANT LAYOUT CONSIDERATION FACTORS The economic construction and efficient operation of a process unit will depend on how well the plant and equipment specified on the process flow-sheet is laid out. Some of the factors considered are:
Cost
Minimization of construction cost is done by adopting shortest run of connecting pipe between equipment. The cost is also reduced by having the least amount of structural steel work. The most important thing is to have an arrangement for best operation and maintenance.
Operation
Equipment such as valves, sample points and instruments are considered as frequently attended equipments. They are located not far away from control room, with convenient positions and heights, to ease the operator’s job. Also, sufficient working and headroom space are provided to allow easy access to equipments.
Maintenance
When laying out the plant, some considerations were made regarding maintenance work. For examples: o Both reactors which use catalysts are located in open space for easiness of removing or replacing the catalysts. o Enough space is allocated for heat exchangers for withdrawing the bundles purposes. 69
o All equipments are accessible to crane/lift truck o Compressors and pumps are located under cover since they require dismantling for maintenance
Safety
Among the safety consideration that we have when laying out this plant are: o Operators have 4 escape routes if anything occurs in the main process unit. o To minimize fire from spread, flammables handling process units are separated from each other o Process vessels with substantial inventories of flammable liquids are located at grade. o Elevated areas will have at least one stairway. o Storage farm which stores the flammable materials are located at safe distance from the main process area. o Equipment subject to explosion hazard is set away from occupied buildings and areas.
Plant expansion
Equipments are arranged by considering future plant expansion, which means it can be conveniently connected with the new equipment. TABLE 29: RECOMMENDED MINIMUM CLEARANCE (SOURCE: PABLO AND MARCELL, 1995)
General Width 30 ft; Headroom 22 ft; distance from
1
Primary roads
2
Secondary Roads
3
Pump Access Aisle Ways
Width 12 ft; headroom 12ft
4
Process Area Main Walkways
Width 10 ft; headroom 8 ft
5
Process Areas Service Walkways
Width 4 ft; headroom 7 ft
6
Main Pipe Racks
Headroom 22 ft
buildings and process area 10 ft Width 20 ft; headroom 20 ft; distance from buildings and process area 5 ft
70
7
Secondary Pipe Racks
Headroom 15 ft The vertical distance between operating
8
Floor Elevations
levels must be no less than height of the tallest process vessel plus 8 ft
Around Hazardous Areas 9
Flare
100 ft
Around Process Equipment Tank Farms: 10 Between tanks
0.5 diameter
11 From tank to dike wall
5 ft
12 Access around diked area
10 ft
13 Dike capacity
largest tank plus 10%
14 Around compressor
10 ft
Between Adjacent Vertical Vessels 15 3 ft diameter
4 ft
16 3-6 ft diameter
6 ft
17 over 6 ft diameter
10 ft
Between Adjacent horizontal Vessels 18 Up to 10 ft diameter
up to 10 ft diameter
19 more than 10 ft diameter
more than 10 ft diameter
20 Between Horizontal Heat Exchangers 4 ft 21 Between Vertical Heat Exchangers
2 ft
71
CHAPTER 6: WASTE TREATMENT 6.1 INTRODUCTION Waste is a general problem in chemical plant operation especially in the developing country where the rules and regulations are very strict regarding the waste disposal. A plant takes few raw materials to produce products through several stages of processes for the sole purpose of generating income. But it is not possible to convert all the raw materials into saleable products thus generating unwanted waste or residual. The wastewater is essentially the water supply of the community after it has been fouled by a variety of usage. Wastewater source of generation may be define as a combination of the liquid or water that carries wastes removed from the residences, institutions and industrial establishments, together with such groundwater, surface water, and storm water.
The
decomposition of untreated accumulation of wastewater will produce large quantities of malodorous gases. It also contains numerous pathogenic microorganisms that inhabit the intestinal tract or that may be present in certain industrial waste. Toxic contaminant in wastewater may lead to fatality of all organisms including aquatic or land inhibited animals and even human beings. Non-biodegradable waste that accumulated in the food chain is absorbed into our body system hence leads to serious sickness such as cancer, food poisoning and others. For these reasons, the treatment of wastewater is necessary in an industrialized society. The wastewater treatment involves the primary treatment for the solid removing, such as screening and sedimentation, the secondary treatment involves biological or chemical treatment and tertiary treatment for the nutrient removal. The increase in environmental awareness has pushed the authority to implement strict regulations to limit the release of proven and potentially hazardous materials by chemical plants.
72
6.2 WASTE MINIMIZATION A waste is best solved from the source. Minimizing the waste is the most effective way to counter the waste problem. This includes reducing or recycling the materials that contribute to waste and resulting in a reduction of total quantity of the waste altogether.
The most preferable way of waste minimization is the prevention of the waste. But this would also be the least economical and logical way of handling the waste because in a process plant, to produce a product, there also will be byproducts created. Byproduct is the main contributor to the unwanted waste.
The other major step in waste minimization is by increasing the efficiency of the process equipment. Increasing the efficiency means less byproduct or waste will be produce by the equipment. But for each percentage of efficiency increment, the tradeoff is the increasing of the overall cost. The cost estimation must be done thoroughly in order to get the optimum operating conditions for all major equipment that contributes to producing waste such as reactor, separator and absorber.
The third option which is highly adapted in all process plant is reuse or recycle of materials. This is the most favorable method in term of cost and the waste minimization.
The least favorable option of waste handling is disposal of the waste. Disposal of the waste to the environment must be under the minimum quantity as possible.
6.3 WASTE AUDIT The discharge from the plant has been identified and there are 3 major wastes. These wastes are divided into three categories; gaseous waste, liquid waste and solid waste.
The gaseous waste is the off gas released from Absorber (C-2). It contains nitrogen, carbon dioxide, carbon monoxide and traces of other materials.
The liquid wastes from Deisobutanizer (C-1) contain isobutane. Besides, other consideration likes dibutyl phthalate discharge or leakage into waste water treatment from Absorber (C-2) and Stripper (C-3). Also from surface runoff during the raining 73
seasons in the plant area contains traces of other materials. Non-acidic waste is from the drainage system in the plant area. Table 30 below shows the streams and type of waste discharged. TABLE 30: WASTE STREAMS PROPERTIES
Parameter
Stream 3
Stream 15
Stream 18
Waste
Isobutane
Off-gas
Water
Type of waste
Liquid
Gaseous
Gaseous
Type of treatment
Recycle and
Recycle,
Recycle to stream 10
storage
storage, and incinerate
Molar flowrate( kmol/hr)
27.50
5798.74
15.27
Temperature (oC)
55.4393
85.7274
85.0
6.4 EFFLUENT DISCHARGE STANDARD AND REQUIREMENTS In Malaysia, Environmental Quality Act 1974 is the act relating to the prevention, abatement, control of pollution and enhancement of the environment, and purposes connected therewith. This parent act, consist of several acts that are enacted from time to time. This plant is subjected to the Environmental Quality (Scheduled Waste) Regulation 2005 which caters for solid waste storage and disposal and Environmental Quality (Sewage and Industrial Effluents)Regulation 1979 which caters for the wastewater released. 6.4.1
Purpose of Effluent Standards
The sole purpose of these Effluent Standards for the discharge from wastewater treatment plants is to control and disposal of effluent to the waters. This will protect the receiving waters and the living aquatic ecosystems. The public health also must be taken into
74
consideration. These standards are crucial because wastewater discharges have been known to contribute considerable amount of the biodegradable organic matter and suspended solids into the receiving waters. These standards stated the maximum values of waste parameters which must not be exceeded in order to release the wastewater into the environment. After taking this into consideration, the design parameters of all the effluent should be less than the standards mentioned in order to ensure that the waste generated by the plant will fall within the required degree. 6.4.2
Liquid waste
There are two standards for effluent discharge specified in the Environmental Quality Act (EQA) 1974: 1. Standard A for discharge upstream of any raw water intake. 2. Standard B for discharge downstream of any raw water intake. The standards are listed in the Third Schedule of the Environmental Quality Act 1974, under the Environmental Quality (Sewage and Industrial Effluents) Regulations, 1979, regulations 8 (1), 8 (2) and 8 (3). An extract of the standards is given below: TABLE 31: ENVIRONMENTAL QUALITY (SEWAGE AND INDUSTRIAL EFFLUENTS) REGULATIONS, 1979 (EXTRACT)
Parameters
Standard A
Standard B
Temperature
40°C
40°C
pH value
6.0-9.00
5.5-9
BOD5
20mg/l
50mg/l
COD
50mg/l
100mg/l
Phenol
0.001mg/l
1.0mg/l
Sulphide
0.50mg/l
0.50mg/l
Oil and Grease
Not Detectable
10mg/l
75
TABLE 32: PLANT WASTEWATER AND STANDARD B VALUES OF EQA
Parameter
Unit
Standard B
Plant wastewater
Temperature
˚C
40
40
5.5-9.0
5.0-9.0
PH value BOD5 at 20˚C
ppm
20
10000
COD
ppm
50
5000
Phenol
mg/l
1.0
>1.0
Oil
mg/l
10
>10
6.4.3
Gaseous Waste
The purge gases from the plant process line include n-butane, nitrogen, oxygen, carbon monoxide, carbon dioxide, water vapor and a little amount of maleic anhydride. The emission of these gaseous is monitored and controlled to meet the requirements of the Environmental Quality (Clean Air) Regulations, 1978 under the Malaysian Clean Air Standards for Dark Smoke and Solid Particle 6.4.4
Solid waste
Under the Environmental Quality (Scheduled Wastes) Regulations 2005, solid waste is categorized under scheduled wastes and must be treated with outmost care. The appropriate safety procedure is required in collecting, packaging, storing and transporting the solid waste. The solid waste is sent to Kualiti Alam Sdn. Bhd. for disposal.
6.5 TREATMENT STRATEGY TABLE 33: METHOD OF REMOVAL ACCORDING TO WASTE COMPONENT
Wastewater Component
Removal Process
Oil and Grease (components of wastewater from other sources)
Aerated Grit removal
Suspended Solid
Screening, Aerated grit removal
Level of Treatment Primary
76
Equalization tank, Clarifier
Volatile Organic Compounds: Maleic anhydride, dibutyl phthalate Acid (H2SO4)and Caustic (NaoH)(used for pH adjustment)
Clarifier, sludge dewatering
Alum Al2(SO4)3 (used as chemical coagulant)
Clarifier, sludge dewatering
Sludge
Sludge dewatering Unit (centrifuge), Mechanical Sludge Dewatering (filter press). Treated sludge shall be sent to Kualiti Alam Sdn. Bhd. for disposal.
Secondary
Tertiary
Besides that, wastes of isobutane from stream 3 (S3) will be recycled and stored in the storage tank for selling.
FIGURE 14: ISOBUTANE STORAGE TANK
77
TREATMENT STAGES PRIMARY TREATMENT 1
SECONDARY TREATMENT 2
3
SCREENING AND
SETTLING AND CHEMICAL
BIOLOGICAL
GRIT REMOVAL
TREATMENT
TREATMENT
7
6
4
NUTRIENT
DISINFECTION
PHYSICAL SEPARATION
REMOVAL
5 SCHEDULED WASTE HANDLING
TERTIARY TREATMENT
78 FIGURE 15: BLOCK DIAGRAM OF WASTEWATER TREATMENT PLANT
Treatment Plant
PRIMARY TREATMENT
SECONDARY TREATMENT
Untreated wastewater
Acid
Alum
Caustic Equalization Tank
Screening
Aerated Grit Removal
pH Stabilizer Clarifier
Sludge dewatering Sludge Cake Mechanical Sludge Thickener
Coagulation Tank
Sludge Two-Stage Centrifugal Separator Spent water
Sludge Storage Area
Disinfection
biopond
For Disposal
FIGURE 16: FLOW SHEET OF WASTEWATER TREATMENT PLANT
TERTIARY TREATMENT 79 TERTIARY TREATMENT
sea
6.6 SCREENING PROCESS This is the first stage of waste treatment which is the primary treatment stage. The screening facility purposes are:
To protect downstream equipments and processes by removing debris and other big particles.
To reduce interference with in plant flow
To minimize blockages in sludge handling facility 6.6.1
Aerated Grit Removal
The grit removal process is also included under the primary treatment stage. The flow velocity is reduced to allow retention time for larger and heavier particles to settle out. The grit removal facility purposes are:
To remove grit that will cause problem to pumps and sludge treatment and dewatering
To remove grease that will cause problem to clarifier
Both grit and grease contain big particles that cannot be broken down by chemical and biological treatment later. 6.6.2
pH Stabilizer
The pH stabilizer facility falls under the secondary treatment which is the chemical treatment stage. It consists of multiple mixers to facilitate the mixing of acid and base in order to stabilize the pH inside the equipment. The purpose of adjusting the pH:
To protect downstream process and equipment from high acidity water that is corrosive in nature.
Non-neutral pH of wastewater is unacceptable at the biological treatment facility because it is toxic to the microorganism 6.6.3
Equalization Tank
Secondary treatment stage also includes the equalization tank. The purposes of equalizing the flow are:
Prevent flow variation for the downstream process
Reduce potential overflow
Reduce hydraulic loading into the downstream process 80
6.6.4
Coagulation Tank
The wastewater must undergo coagulation process before entering the tertiary treatment which is the biological treatment. The purpose of coagulation process:
Removal of suspended and colloidal solid which cannot be removed by sedimentation.
Reduce soluble organic content in the wastewater consequently reducing the COD and BOD values.
Removal of metals, phosphorous and colored substances 6.6.5
Clarifier
The first phase of tertiary treatment stage is the clarifier facility. It provided the sedimentation time which reduces the velocity of wastewater that will allow organic matter in suspended solid to settle out. The purpose of this facility:
Remove maximum amount of solids
Separate wastewater into sludge and spent water which will be treated separately for optimum efficiency. 6.6.6
Sludge Dewatering
All treatment processes will generate significant amount of sludge containing inert and nonbiodegradable organic matter. This particular type of sludge must be dispose because it cannot be treated anymore. Purpose of sludge dewatering facility:
Remove the water contain in the sludge.
Separate the water from sludge
Turned the sludge into sludge cake with low percentage of water. 6.6.7
Mechanical Sludge Thickener
The main purpose of this equipment is to thicken the sludge cake from 1% dry solid to about 6% dry solid content. The thickening equipment used is the gravity belt thickener. To increase the thickening process, a chemical dosing conditioning also injected into the sludge. 6.6.8
Sludge Storage
The dried sludge cake will be stored for 30 days before disposal to allow sufficient accumulation of the required quantity of sludge to be disposed. The storage building or 81
structure must have a roof with partly open walls to allow good ventilation. For good conduct and safety, the storage area should be situated downwind. 6.6.9
Disinfection
The main reason for the disinfection facility is to destruct selective disease causing organism in the wastewater. Disinfection is important for the wastewater that will be released out into the open water system. Usage of calcium hypoclorite is the most preferable option because the typical chlorination type of disinfection is very harmful to operator. Good mixing during the disinfection stage is important. 6.6.10 Biopond The biopond facility will provide the last biological treatment before releasing the wastewater to sea. The wastewater from the disinfection phase will be exposed to microorganism that will dissolve the remaining organic substances. But the ratio of wastewater from the disinfection phase must be kept very low to the ratio of wastewater contained in the biopond. This is to facilitate the dilution process where all the traces of calcium hypoclorite would not have any affect anymore. This is crucial to prevent toxicity to microorganism in the biopond. The holding period of treated wastewater in the biopond will range from one to two week before releasing to receiving water. So the volume of the biopond will have to be bigger.
6.7 GAS TREATMENT When making contact with the gaseous waste, there are few methods available for the treatment. The methods are incineration, condensation, adsorption and flaring. All the treatment method has been studied and the limitation of each method is taken into account when choosing the best method to be used to treat the gaseous waste. The limitation as shown in the table: TABLE 34: SUMMARY OF THE LIMITATIONS FOR THE GASEOUS TREATMENT STRATEGY
Treatment
Limitation
Method
More economical since the gaseous wastes is not going to be recover
Control by
or as the result of intermittent, uncertain or emergency process
flaring
operations
The combustion of VOCs will produce harmless or much less 82
harmful substances since the flare temperature will be operating below 1000 K to avoid the formation of NOx
Flare will be injected with steam to enhance mixing so that the combustion process will be as complete as practical
Normally used for large VOCs content stream for recovery
Control by
Not economical for small stream
adsorption
Activated carbon is a very effective adsorbent in removing VOCs, but quite expensive
Not practical for low gas flow
The N2 presents in the gas stream may enter the atmosphere partly as N2, NO or NO2
Incomplete combustion of gas stream can produce an exhaust gas that is more harmful than the input gas
Control by
incineration
Additional fuel is require to burn the VOCs if the total mass fraction of VOCs too small
If a heat exchanger is installed to lower the cost of fuel, the cost of the heat exchanger itself is high and may lead to corrosion problem
If catalytic incinerator is applied, the fuel cost is greatly reduced and the operating temperature is low. However, the catalyst will significantly increase the operating cost
Normally used for large VOCs content stream for recovery purpose
This method is not economical for small stream
The temperature is low enough that ordinary one-stage refrigerators cannot be used
Control by
Often the temperatures required for high removal efficiency are below the freezing temperature of the material being removed so that
condensation
the material freezes on the cooling coils, requiring frequent defrosting
If the gas being treated contains significant amounts of water vapor, it will condense and freeze on the cooling coil, this requiring frequent defrosting
83
After considering all the options, the flaring method has been chosen because it fulfills the entire requirement and suitable to treat the gaseous waste discharge at the off gas line. Its systems are the satisfactory provision in the design and operation of the basic requirements for combustion. These gases must be release to maintain the operating pressure of the equipment in the process plant from daily operation and also during the emergency shutdown. The main control that needs to be maintained along the flaring process is the control of proper steam flow. This is because with proper steam flow, smokeless operation can be maintained at all conditions of gas flow, which provide an almost complete combustion of gaseous. To conform to the Environmental Quality (Clean Air) Regulations 1978, a filter should be installed at the stack gas tip before releasing the gas. Gas quality monitoring system should also be installed in order to ensure that the gas release is within the acceptable range. Besides that, the nitrogen gaseous that was produced at stream 15 (S15) will be separated by using nitrogen separator and will be used as process blanket. The other component such as isobutane will be recycled and kept in the storage tank for selling.
FIGURE 17: NITROGEN SEPARATOR
For the gas waste from the stream 18 (S18) will also be channeled to the flare for combustion.
84
FIGURE 18: WASTE OF STREAM 18 (S18)
6.8 SOLID HANDLING TREATMENT Solid wastes of the plant are generated from different processes, chemical handling operations as well as from wastewater treatment, which consists of hazardous and nonhazardous wastes. Hazardous waste is defined as any solid waste listed as hazardous or possesses one or more hazardous characteristics as defined in federal waste regulations. Its effect can last for very long periods of time. Major solid wastes are typically in the form of sludge, scrap and spent process catalyst and it’s divided into three; scheduled waste, recycled waste and domestic waste. There are a few treatment methods in handling the solid waste and it have been summarized in Table 35 below: TABLE 35: DISPOSAL METHODS OF SOLID WASTE
Solid Waste
Disposal/Treatment Method
Tank bottom sludge (from Wastewater Sent to KualitiAlam Plant) Empty Drums
Returned to supplier
Wood, metal scrap/various valuables like
Sold to contractors
empty drums Miscellaneous (paper, plastic, domestic
Disposed off through a Contractor
waste)
85
6.9 SCHEDULED WASTE Scheduled wastes, one of the solid wastes produced in the plant is a small percentage of hazardous waste, which has been regarded for a long time as intractable, or difficult to safely dispose of, without special technologies and facilities. It is also can be defined as a material or article containing a chemical, or mixture of chemicals, exceeding the threshold concentration and threshold quantity which are:
organic in nature
resistant to degradation by chemical, physical or biological means
toxic to humans, animals, vegetation or aquatic life
bio-accumulative in humans, flora and fauna
According to Environmental Quality (Scheduled Wastes) Regulations, 1989, solid waste is categorized under scheduled wastes, and must be treated appropriately. It is the duty of the plant management to adopt safety procedure in collecting, packaging, inventorying and transporting the solid waste to suitable parties before further treatment. Sludge is turned into sludge cake to reduce the weight and for ease of loading into plastic lined drums before being transported.
86
CHAPTER 7: PROJECT ECONOMICS AND COST ESTIMATION 7.1 INTRODUCTION The economic evaluation of a plant is important in determining the profitability of a plant in generating profit. Therefore, before building any plant, the design engineer needs to decide between alternative designs to be implemented as well as the overall plant economics. In evaluating the project economics, estimates of investment and equipment costs are required. Before the final process design starts, company management normally becomes involved to decide if significant capital funds will be committed to the project or not. 7.1.1
Capital Investment
The estimation of Total Capital Investment and Total Product Cost of the project are determined by using the methods suggested by Peters and Timmerhaus .Equipment purchasing amount are determined by using method by Warren D.Seider, J.D Seader and Daniel R.Lewin. The capital needed to supply the required manufacturing and plant facilities is called fixed capital investment (FCI) while that necessary for the operation of plant is termed the working capital (WC). Start-up Cost (SC) is the cost required at first once the process of the plant is started. Thus the sum of the fixed capital investment, working capital start-up cost is known as the total capital investment (TCI). Furthermore, cash flow and discounted cash flow are also constructed in determining the Pay-Back Period as well as Net Present Value for the project. Fixed capital investment can be divided into two that are manufacturing fixed capital investment (direct cost) and non-manufacturing fixed capital investment (indirect cost). Basically, FCI depends on the total equipments cost available in the plant multiply with a factor that varies according to what type of cost it represents. Meanwhile, working capital is the additional investment needed, over and above the fixed capital, to start the plant up and operate it to the point when income is earned. The calculation made follows the Douglas’s approach method The interest rate for plant operation is 10% per annum Project life-cycle will be 15 years Plant operates at normal annual operation period which is 330 days 87
7.1.2
Total Equipment Cost (TEC)
Capital cost estimate for chemical process plants can be based on purchase cost estimation of the major equipment items. The equipment cost will be used along with factors for estimating other relevant costs(Sinnott.2005). TABLE 36: TOTAL EQUIPMENT COST
Expander
1000000
1
1000000
Total Cost (RM) 3059039.462
Centrifugal Separator
1340000
1
1340000
4099112.879
160000
1
160000
489446.3139
Cooler 1
59500
1
59500
182012.848
Cooler 2
67000
1
67000
204955.6439
Heat Exchanger 1
245000
1
245000
749464.6681
Heat Exchanger 2
116000
1
116000
354848.5775
Heat Exchanger 3
134500
1
134500
411440.8076
1580000
1
1580000
4833282.349
Recycle Pump
13000
3
39000
119302.539
Deisobutanizer
570000
1
570000
1743652.493
Stripper
343000
1
343000
1049250.535
MAN Tank
238000
1
238000
728051.3919
Reaction Reactor (incl catalyst)
620000
1
620000
1896604.466
Mixer
160000
1
160000
489446.3139
Equipments
Butane Tank
Absorber
Cost
Quantity
TOTAL EQUIPMENT COST
Total Cost (USD)
20409911.29
In order to estimate the capital cost for chemical process plant, the factorial method of cost estimation is used. To make a more accurate estimate, the cost factors that are compounded into the ‘Lang factor’ are considered individually. 7.1.3
Fixed Capital Investment
The direct cost and indirect cost items incurred in the construction that are used to calculate Fixed Capital Investment of a plant are as such : FCI
= Total Equipment Cost * Lang’s Factor = RM 71965347.20 (in 2004) 88
Fixed Capital Investment (FCI)
= FCI (2004) * (CEPCI 2012/CEPCI2004)
= 94889923.13 Total Capital Investment = Working Capital (15% FCI) + S/Up Cost (10% FCI) + FCI = RM 118 612 403.90 7.1.4
Estimation of Total Operating Cost
All expenses directly connected with the manufacturing operation or the physical equipment of a process plant is included in the operating costs. Raw Material Utilities Maintenance Cost Operating Supplies Operating Labour Management Personnel Laboratory Charges Patent and Royalties TOTAL VARIABLE COST
Cold Utilities 2% of FCI 10% of Maintenance Cost 10% of Operating Labour 5% of Operating Labour 1% of Total Expenses
98838109.36 25192299 1897798.463 189779.8463 720000 72000 36000 1269459.867 128215446.5
Local tax Insurance FIXED CHARGES
1% of FCI 0.4% of FCI
948899.2313 379559.6925 1328458.924
PLANT OVERHEAD TOTAL MANUFACTURING COST
50% (Maintenance + Labour + Supervision)
1344899.231 130888804.7
Administrative Cost Distribution and Selling R&D TOTAL GENERAL EXPENSES
15% (Maintenance + Labour + Supervision) 10% of TMC 5% of TMC
403469.7694 13088880.47 6544440.235 20036790.47
TOTAL OPERATING COST = Total Manufacturing Cost + General Expenses = RM 150,925,595.20 7.1.5
Gross Profit
For the case of this plant, the revenue comes from solely selling of MAN as shown below: Profit = Total Sale – Total Operating Cost = (RM 7 944.33/ton x 50000tonne/yr) – 150,925,595.20 = RM 177,395,904.80/year 89
7.2 PROFITABILITY ANALYSIS In determining the economic attractiveness of a project, it is important to based decision three important economic parameters which are Investment Rate of Return (IRR), Net Present Value (NPV) and Pay Back Period (PBP). Several assumptions are made in the economic analysis of this maleic anhydride acid plant. The assumptions are as follows: The plant has a project plant life of 15 years. The plant construction period is 3 years before commencing production. Hence, the total investment cost is distributed evenly between the 3 years. Interest rate is 15% Local Taxes is assumed to be at 10% annually. 7.2.1
Start-up Period
The plant will start up in year 4. The plant construction assumed to be last for 3 years. TABLE 37: CAPITAL INVESTMENT FOR START-UP PERIOD
Start up period of 3 years
Year 0
Equipment & Design
- 29602641.93
Year 1
(Direct Cost – Year 1 Cost)/2
-19895528
Year 2
As year 1
-19895528
Year 3
Indirect Cost + S/up cost
-26294129.60
Total
7.2.2
-95687827.53
Depreciation
When government taxing comes into place, depreciation becomes important to aid the plant from balancing tax payment to equipment wear. In this project, depreciation per annum can be computed as.
B SV (B - 0.03 Fixed Capital Investment) n 15 RM 118612403.9 (0.03 RM 94889923.13) 15 RM 7.7million Dt
90
where t year (t 1,2......, n) D t annual depreciation charge B total investment SV estimated salvage value, 0.03x Fixed Capital Investment n expected depreciable life 7.2.3
Cash Flow Estimation
Here, the net cash flow in each year of the project is brought to its “present worth” at the start of the project by discounting it at some chosen compound interest. n t
NFW n n 1 ( 1 r)
Total NPW of projec t
where, NFW estimated net cash flow in year n r discount rate(interest rate) per cent/100 t
life of projects, years
A discount rate of 15% is used From the investment as Table 37 and also annual gross earning, cash flow diagram is as follow, to obtain the payback period.
FIGURE 19: CASH FLOW DIAGRAM
91
7.2.4
Net Present Worth TABLE 38: CUMULATIVE CASH FLOW
r (%) NPW (Cumulative Cash Flow at years 15) (RM) 10 20 30 40 50
630341893 255582162 102887451 33344914 -1062648
From the net present worth calculated by varying the discount rate, it will reach negative when the rate is in between 40-50%, this is higher than MARR, thus project is feasible. 7.2.5
Internal Rate of return
7.2.6
Rate of Return (ROR) Estimation
The simplest method is to base the ROR on the average income over the life of the project and the original investment. Cumulative net cash flow at en d of project 100 per cent Life of pr oject original i nvestment 399036617 100% 22.43% 15 118612403.91
ROR
Expected ROR (Rate of Return) must at least meet or exceed the MARR (Minimum Attractive Rate of Return) of 15%. Since, the rate of return of the project is 22.43%, which is more than 15%, the project is worth investing. 7.2.7
Net Present Value or Worth (NPV) Estimation n t
NFW ' n n 1 ( 1 r )
NPW where,
NFW estimated net cash flow in year n r discounted cash flow rate of return per cent/100 t
life of projects, years
NPV calculated is RM 399 million
92
7.2.8
Pay Back Time
Pay Back Period the time that must elapse after startup until cumulative undiscounted cash flow repays fixed capital investment. For this project, it is 5 years.
Payback Period Cumulative Cash Flow (RM)
2E+09 1.5E+09 1E+09 500000000 0 0
5
10
-5E+08
15
Year FIGURE 20: PAYBACK TIME
7.3 DISCUSSIONS From the economic analysis done, the plant will need a capital investment of RM RM118.612 million. An internal rate of return (IRR) of 49 % and a net present value (NPW) of RM 399 million is attainable for a project plant life of 15 years with a payback time of 5 years after plant startup. In addition, annual sales will generate an income of RM 177.395million annually. In fact the return on investment is 22.43%, which is higher than MARR, hence, this indicates that the plant is economically viable and economically attractive. The economic evaluation conducted on the maleic anhydride plant at this stage serves only as a very crude estimation. Furthermore, the correlations taken from different references may give different value and methods in estimating the plants cost. In addition, the assumptions made may also be invalid, considering the fact that economic environment is very sensitive to the global environment and may subject to changes from time to time. Apart from that, it is very difficult to predict the actual annual cash flow since the reliability of the plant equipments is not known. Hence, the actual investment cost may be larger than the predicted figure. For improvement, a concise economic evaluation should be carry out by considering all the factors mentioned above to obtain the best economic potential for the MAN plant.
93
CHAPTER 8: CONCLUSION & RECOMMENDATION
8.1 CONCLUSION The overall design project for the production of Maleic Anhydride meets the desired requirements and objectives. The production of 50,000 metric tonne per year of Maleic Anhydride has a bright future especially in the South East Asia region. From the feasibility study carried out, the future growth of Maleic Anhydride demand in the South East Asia market is optimistic, 4.5 % per year growth on Maleic Anhydride consumption is forecasted until year 2009. The proposed plant is situated at Kidurong Industrial Estate which is feasible, economically and environmentally. Moreover, the utilities supply such as cooling water supply, deionized water, electricity and steam supply required for the process can be easily obtained. The process chosen for the production of Maleic Anhydride is the catalytic oxidization of nbutane. Heuristics approach has been applied in identifying the appropriate design. For the process synthesis and flow sheeting, base case material and energy balance together with process simulation has been performed both by manual calculation and iCON simulator. Detailed equipment process design and mechanical engineering design of all major equipment has been performed. A highly integrated heat exchanger network and process control system is also included to the proposed plant to ensure the profitability of the plant. In responding to the environmental responsibility, the plant has been designed to achieve the target of waste minimization and cost minimization. The unwanted side product is being combusted and well treated to ensure the emission coming out from the plant has met the Malaysian government Environment Quality Act, 1974. On the safety aspect, HAZOP study has been conducted to identify the occurrence of operational problems and providing necessary resolution. A general safety study includes personal safety, emergency management, Standard Operation Procedures (SOP) and plant start up and shut down procedures were documented in this report as well.
94
As for the economic evaluation of the process plant, the cost of the plant is calculated by using detailed factorial method. The total capital investment of the plant is approximately RM 118 million. The Rate of Return (ROR) is 22.43 % which is higher than Minimum Attractive Rate of Return (MARR) 15 %, therefore the project is worth investing. The payback period is 5 years after plant start-up which is within feasible period of time. Finally, it can be concluded that the construction of a 50,000 metric tonne per year of Maleic Anhydride production plant in Kidurong, Bintulu is technically feasible and economically attractive.
8.2 RECOMMENDATION The final year design project has been useful in cultivating and enhancing the skills and knowledge at hand. As final year students, the experience gained throughout the process of this project has given the opportunity to actually design a real processing plant has increase our understanding in the chemical engineering field. Besides that, other skills were also developed in the process such as communication skills, management skill, and team work. However, we find that there few areas that needs improvements and perhaps a new approach in solving some of the problems faced. Firstly, we recommend that the PDP committee provide standardized values of key elements such as the feeds prices. We have different values of feed because taking it from different sources. So, a standardized value is appropriate to accommodate the student with a good reference. Secondly, further research must be done on isomeric conversion technology in order to convert isomers back to its normal state. This is important to reduce the amount of byproduct produced from the process. Thirdly, we propose that students should be provided easier access to the labs to use engineering software such as HYSYS and AutoCAD and the department should provide a proper manual as guidance. Finally, a complete and clear guidance should be provided to the student and without further amendment. Students are facing difficulty to cope with sudden changes and added requirements made by the coordinator due to the short time frame to finish it. We hope that these recommendations will be considered and useful by the PDP committee to improve the PDP project handled by the student in the future.
95
REFERENCES (1992). Maleic Anhydride World Survey. London: Tecnon(U.K.) Ltd. Maleic Anhydride. (2011, November). Retrieved June 2, 2012, from IHS Chemical: http://www.ihs.com/products/chemical/planning/ceh/maleic-anyhydride.aspx Maleic Anhydride. (2011). Retrieved June 4, 2012, from Thirumalai Chemicals Ltd.: http://www.thirumalaichemicals.com/maleic.html AP-42,
CH
6.14:
Maleic
Anhydride.
(n.d.).
Retrieved
June
1,
2012,
from
www.epa.gov/ttnchie1/ap42/ch06/final/c06s14.pdf BDA. (2010, April 10). Official Website of Bintulu Development Authority. Retrieved June 11,
2012,
from
Bintulu
Development
Authority
Official
Website:
http://www.bda.gov.my/modules/web/index.php?menu_id=0&sub_id=1 Fogler, H. S. (2006). Elements of Chemical Reaction Engineering (4th ed.). New Jersey: Pearson Education Inc. Ishak, M. (2011, September 1). Bintulu Analysis. The Report, SARAWAK 2011, 34-43. Jose, S. (2008, November 3). World Maleic Anhydride Market to Reach 2.0 Million Metric Tons by 2012, According to New Report by Global Industry Analysts. Retrieved June 4,
2012,
from
PRWeb:
http://www.prweb.com/releases/maleic_anhydride/butanediol/prweb1553754.htm KKIP. (1995, November 2). Investing in KKIP. Retrieved June 12, 2012, from Kota Kinabalu Industrial Park: http://www.sabah.com.my/kkip/inv.html Lohbeck, K., Haferkorn, H., Fuhrmann, W., & Fedtke, N. (2005). Maleic and Fumaric Acids. In Ullmann's Encyclopedia of Industrial Chemistry. Weinheim: Wiley-VCH. Pinch Analysis. (n.d.). Retrieved August 25, 2012, from KBC Nextgen Performance: http://www.kbcat.com/?id=51&fm=search&searchText=pinch%20analysis Production
of
Maleic
Anhydride.
(n.d.).
Retrieved
June
6,
2012,
from
www.che.cemr.wvu.edu/publications/projects/large.../maleic.PDF Production of Phthalate Anhydride from O-xylene. (n.d.). Retrieved June 7, 2012, from www.che.cemr.wvu.edu/publications/.../phthalic2/phthalic2-b.pdf 96
Sarawak Government. (2010, April 20). Sarawak Government Portal | Investment Incentives. Retrieved June 11, 2012, from The Official Portal of Sarawak Government: http://www.sarawak.gov.my/en/investors/investment-incentives Silla, H. (2003). Chemical Process Engineering Design and Economics. USA: Marcel Dekker. Timbang, M. (2007, March 2). Bintulu - Wikipedia, the free encyclopedia. Retrieved June 11, 2012, from Wikipedia: http://en.wikipedia.org/wiki/Bintulu Timothy R. Felthouse, J. C.-J. (2001, April 26). Department of Chemistry. Retrieved June 2, 2012,
from
University
of
South
Alabama:
http://www.southalabama.edu/chemistry/barletta/felthouse.pdf Tissue, B. M. (2000). The Chemistry Hypermedia Project. Retrieved August 1, 2012, from CHP website: http://www.files.chem.vt.edu/chem-ed/sep/gc/gc.html Town and Regional Planning Department Sabah. (2011, January 1). Land Zoned for Industry Sabah. Retrieved June 12, 2012, from Jabatan Perancang Wilayah dan Negeri Sabah: http://www.townplanning.sabah.gov.my/ US EPA. (2011, July 8). Oil and Gas Production Waste | Radiation Protection | US EPA. Retrieved
July
25,
2012,
from
US
Environmental
Protection
Agency:
http://www.epa.gov/rpdweb00/tenorm/oilandgas.html#scale Woril Turner Dudley, V. K. (2012, January 3). Maleic Anhydride - Process Design. Retrieved June 2, 2012, from Scribd.: http://www.scribd.com/doc/76994917/MaleicAnhydride-Process-Design#
97
APPENDICES APPENDIX 1: PROCESS FLOW DIAGRAM OF MAN PLANT .................................................I APPENDIX 2: P&ID OF MAN PRODUCTION PLANT ............................................................ II APPENDIX 3: CRITERIAS FOR EACH SITE PLANNED ...................................................... III APPENDIX 4: MASS BALANCE CALCULATION (DEISOBUTANIZER)............................... V APPENDIX 5: MASS ENERGY BALANCE (MIXER) ............................................................. IX APPENDIX 6: MASS ENERGY BALANCE (REACTOR) ......................................................... X APPENDIX 7: MASS ENERGY BALANCE (ABSORBER) ..................................................... XII APPENDIX 8: MASS ENERGY BALANCE (STRIPPER) ..................................................... XIII APPENDIX 9: CP VALUE FOR ENERGY BALANCE CALCULATION ............................ XVII APPENDIX 10: DRAFT OF PLANT LAYOUT ..................................................................... XIX
98
APPENDIX 1: PROCESS FLOW DIAGRAM OF MAN PLANT
i
APPENDIX 2: P&ID OF MAN PRODUCTION PLANT
ii
APPENDIX 3: CRITERIAS FOR EACH SITE PLANNED
Kidurong Industrial Area
Kota Kinabalu Industrial Park
PasirGudang Industrial Estate
Location
20 km from Bintulu Town
25 km from KK
36km from Johor Bharu
Type of industry
Light & Medium
Any compatible
Light, Medium & Heavy
Petrochemical and gas Timber-based Plantation and Agro Energy Intensive 97.3 hectare RM77.42
Food Timber-based Plantation-based
Selection Criteria
Preferred Area available Land price(per m2) Raw material Supplier Power Supply
SESCO’S Combined Cycle Power Plant (132MW) Bakun Hydroelectricity Power Project (2400MW) Sarawak Power Generation Plant (220MW) Bintulu Water Supply Treatment Plant
7.05 acres RM129.17
430 acres RM 86.08 – 236.72 Optimal, Kerteh Amoco Chemicals, Gebeng
KKIP Power SdnBhd
Sultan Iskandar Power Station (644 MW)
(300MW) Powertron Resources S/B
IPP YTL Power Generation Sdn. Bhd.
(120MW) Sabah Electricity SdbBhd (293MW)
100 acres RM20 PETRONAS Gas Berhad TasikKenyir Hydroelectric Dam (400MW) IPP YTL
(600 MW)
Paka Power plant CUF Kerteh
Diversified Water Resources S/B
Bintulu Deepwater Port
Sepanggar Bay Port
Airport
Bintulu Airport
Kota Kinabalu International Airport
Railway facilities
-
-
PengerangIntergrated Petroleum Complex 42km from Johor Bharu Petrochemical and refinery
Petrochemical
Water Supply
Port Facilities
Kertih Integrated Petrochemical Complex 5 km from Paka and 9.6 km from Kemaman Petrochemical, Chemical and General
iii
22, 500 acres RM64.58-RM86.11 PETRONAS Rapid (LNG Regasification Plant) Sultan Iskandar Power Station (TNB) PasirGudang Power Station (YTL Power International Bhd) PETRONAS Power Plant (future planning)
Loji Air Sungai Layang
Bukit Sah
SAJ Holding SdnBhd
Syarikat Air Johor
Sungai Cherol
Future investment on water supply project
Loji Air Sungai Buluh
Sungai Kemasik
PasirGudang Port
Kerteh Port, TanjungBerhala Port, Kuala Terengganu port
Senai International Airport, Johor and Changi International Airport, Singapore Singapore and North
Sultan Mahmud Airport, Kuala Terengganu and Sultan Ahmad Shah Airport, Kuantan Kuantan-Kerteh Railway
Pengerang Petroleum Terminal TanjungLangsat Petroleum Terminal Senai International Airport, Johor and Changi International Airport, Singapore KTM Singapore-North
Peninsular Malaysia Pan-Borneo Highway Roadways
Incentives
peninsular route
KK-Sulaman Road
Main road to Singapore
Kuala Terengganu-Kuantan-
Second Link Expressway
KK West Coast Parkway
PLUS Highway
Karak Highway
Senai-Desaru Expressway
Pioneer Status 5-years 70% tax exemption on statutory income Investment Tax Allowance Allowance of 100% in respect of qualifying capital expenditure incurred Reinvestment Allowance Allowance of 100% in respect of qualifying capital expenditure incurred
Federal Route 500 Pioneer Status 5-years corporate tax on 15% of statutory income Investment Tax Allowance Allowance of 85% in respect of qualifying capital expenditure incurred Infrastructure Allowance 100% infrastructure allowance on qualifying expenditures
Incentives for High Tech Industries
Land Incentives by State Government
Rebate on industrial land
Free Infrastructures
North-South Highway Infrastructure Allowance.
a flat corporate tax rate of 3% of chargeable income;
Incentives for research development
Five-year exemption on import duty.
100% exemption on director fees paid to non-Malaysian director;
Exemption from import duty on direct raw materials/components
5 % discount on monthly electrical bills for first 2 years.
50% exemption on gross employment income for nonMalaysian professional traders;
Pioneer Status and Investment Tax Allowance and Reinvestment Allowance. Incentives for high tech industries and for training tariff protection
25-38 % exemption on daily water cost for 4545 m3 of water for Pioneer Status and Investment Tax Allowance and Reinvestment Allowance.
Incentive for exports
tax exemption of stamp duties on documentation tax exemption on dividends received by or from the LITC companies.
Incentives for high tech industries Local people (Below 40 years old)
200 000 peoples
200 000 peoples
1500 000 peoples
200 000 peoples
40,000 people
Effluent Treatment Plant of CUF KualitiAlamSdnBhd, Bukit Nenas, Negeri Sembilan
Waste water management
KualitiAlamSdnBhd Indah Water Konsortium
iv
KualitiAlamSdnBhd, Bukit Nenas, Negeri Sembilan
APPENDIX 4: MASS BALANCE CALCULATION (DEISOBUTANIZER)
Example Mass Balance Calculation for Deisobutanizer distillation column Calculation of component distribution using Hengstebeck Method
d log i m log ij C bi Calculation of saturation pressure of each component using the equation below:
The operating condition of the column is 850C and 11atm. The light key (LK) is isobutene distilled in the top at 55% and the heavy key (HK) is nbutane distilled at the bottom at 99.9% The feed flow rate is 9700 kg/h with the pre-specified feed composition. Example calculation using propane:
Partial pressure, Pi = Psat(x) = 3487.14(0.02) = 69.74 kPa Volatility, Ki = Pi/PT = 69.74/1114.575 = 0.063 Relative volatility, α = Ki/Khk = 0.063/0.7 = 0.09
v
The calculation of other component is done below: MW
Component
Mass
Mass flow rate
Mole flow rate
Mole Fraction
Psat
Pi
Ki
fraction 44.096
Propane
0.0154
149.38
3.3876089
0.02023
3487.1483
70.54017
0.06328885
58.122
Isobutene
0.295
2861.5
49.232649
0.29399
1535.3106
451.35917
0.404960784
58.122
n-butane
0.677
6566.9
112.98476
0.67467
1162.3554
784.20863
0.703594309
56.106
Isobutene
0.0013
12.61
0.2247531
0.00134
1404.7017
1.8852229
0.001691428
56.106
1-butene
0.002
19.4
0.3457741
0.00206
1367.9095
2.8243767
0.002534039
72.149
Neopentane
0.0011
10.67
0.1478884
0.00088
881.58433
0.7785218
0.000698492
72.149
iso-pentane
0.0077
74.69
1.0352188
0.00618
524.91238
3.2448288
0.00291127
72.149
n-pentane
0.0008
7.76
0.1075552
0.00064
426.31218
0.2737991
0.000245653
Α
log α
Propane
0.08995077
-1.0459951
Isobutene
0.57556006
-0.2399093
n-butane
1
Isobutene
0.00240398
-2.6190689
1-butene
0.00360156
-2.443509
Neopentane
0.00099275
-3.0031608
iso-pentane
0.00413771
-2.3832398
n-pentane
0.00034914
-3.4569996
Component
0
vi
A graph of log d/b and log α is plotted for the LK and HK as reference for finding the composition of the other components. Example calculation for isobutene (LK): d = 0.55(49.23) = 27.08 b = 0.45(49.23) = 22.15 d/b = 27.08/22.15 = 1.22 log d/b = 0.087 log α = -0.24 Component LK
HK
F
49.23265 112.9847562
D
27.07796 0.112984756
B
22.15469 112.8717714
d/b
1.222222 0.001001001
lg d/b
0.08715
-2.99956549
lg α
-0.23991
0
log d/b vs log α 0.5 0 -0.25
-0.2
-0.15
log d/b
-0.3
-0.1
-0.05
y = -12.866x - 2.9996
-0.5 0 -1 -1.5 -2 -2.5 -3
log α
From the graph above, m = -12.866 and c = -2.9996. The d/b for other components are calculated. Example calculation for propane: vii
-3.5
Mol fraction going to bottom from feed, xbi = 1/(d/b+1) = 0 Mol fraction going to distillate from feed, xdi = 1- xbi = 1 Mol flow rate component in bottom, L = xbi(mol flow rate component) = 0 Mol flow rate component in distillate, V = xdi(mol flow rate component) = 1(3.388) = 3.388 mol/hr Component propane isobutane n-butane isobutene 1-butene neopentane iso-pentane n-pentane
log α
log di/bi
di/bi
Xbi
xdi
V
L
-1.0459951
10.4581731
2.8719E+10
3.482E-11
1
3.38760885
1.1796E-10
-0.2399093
0.08707368
1.22200697
0.45004359
0.54995641
27.0758105
22.156838
0
-2.9996
0.00100092
0.99900008
0.00099992
0.11297579
112.87178
-2.6190689
30.6973407
4.9813E+30
2.0075E-31
1
0.22475315
4.512E-32
-2.443509
28.4385868
2.7453E+28
3.6426E-29
1
0.34577407
1.2595E-29
-3.0031608
35.6390672
4.3558E+35
2.2958E-36
1
0.1478884
3.3952E-37
-2.3832398
27.6631637
4.6043E+27
2.1719E-28
1
1.03521878
2.2484E-28
-3.4569996
41.4781574
3.0072E+41
3.3254E-42
1
0.1075552
3.5766E-43
viii
APPENDIX 5: MASS ENERGY BALANCE (MIXER)
Example Mass Balance Calculation for mixer
4 7 6 Input = output Stream 4 + stream 6 = stream 7 Component
Steam 4
Steam 6
Steam 7
(kg/h)
(kg/h)
(kg/h)
Isobutane
1287.80
0
1287.80
n-butane
6560.33
0
6560.33
Oxygen
0.00
39730.8667 39730.87
Water
0.00
3195.31948 3195.32
Nitrogen
0.00
130843.358 130843.36
Stream 6 is the air feed with a 65% humidity. Thus, from the Psychometer chart from appendix figure 5.1 H=0.018kg water/kg dry air =0.03 kgmol water/kgmol dry air The oxygen feed must be below 0.0181 mol % which is below the flammability limit of nbutane mixture. By using excel, the required amount of oxygen is 39730 kg/h Total Nitrogen = 39730(0.79/0.21) = 130843.358 kg/h Total air = 42926.187 kg/h = 5912.33 kgmole/h Total water with the air = 5912.33(0.03) = 177.37 kgmole/h = 130843.358 kg/hr
ix
APPENDIX 6: MASS ENERGY BALANCE (REACTOR)
Example Mass Balance Calculation for Reactor C4H10 + 3.5O2 C4H2O3 + 4H2O C4H10 + 5.5O2 2CO + 2CO2 + 5H2O Using extent of reaction method, with the assumption of pure n-butane feed and oxygen feed is used for the sake of simplicity. The actual process uses a mixture of n-butane feed with air at 65% humidity used as the source of oxygen. The water entering together with air is calculated and the ratio of air to feed is kept at a ratio of 0.017% which is below the flammability limit.
The following initial are used for reference: b = n-butane, f = final, i = initial, MAN = maleic anhydride, CO = carbon monoxide, CO2 = carbon dioxide, O2 = oxygen
Production: No. of moles for n-butane:
No. of moles for MAN:
No. of moles for carbon monoxide:
No. of moles for carbon dioxide:
No. of moles for water:
No. of moles of oxygen:
Conversion = 0.822 Selectivity of MAN; Selectivity of CO; x
Selectivity ratio of MAN/CO;
Assuming,
Conversion of n-butane:
Consumption of oxygen:
Production:
n-butane = 8081.283 kg/h Oxygen = 13715.04 kg/h
9 Reactor
8
xi
n-butane = 1438.52 kg/h Maleic Anhydride = 9805.7 kg/h Carbon Monoxide = 800.526 kg/h Carbon Dioxide = 1257.806 kg/h Water = 8750.606 kg/h
APPENDIX 7: MASS ENERGY BALANCE (ABSORBER)
Example Mass Balance Calculation for Absorber column According to the US patent 4118403, the dibutyl phthalate (DBP) absorbs 99.4% of the MAN and 0.1% of the water in the stream. The gas class components are assumed to exit as gases at the top of the absorber. Volume feed entering absorber = 100217.558 m3/h Amount of dibutyl phthalate needed according to US patent 4118403 is 0.2 kg DBP/ m3 feed entering. Amount DBP needed = 0.02(100217.558) = 20043.51 ≈ 20000kg/h At the bottom of absorber: Amount MAN = 0.96(7518.16) = 7086.58 kg/h Amount water = 0.001(10021.49) = 100.21
Isobutane = 1287.80 kg/h n-butane = 1167.74 kg/h Oxygen = 27884.01 kg/h Maleic anhydride = 71.58 kg/h Carbon monoxide = 1168.42 kg/h Carbon dioxide = 1835.85 kg/h 11 Water = 9921.27 kg/h Nitrogen = 130843.36 kg/h Dibutyl phthalate = 20000kg/h 17
10
Absorber
Isobutane = 1287.80 kg/h n-butane =1167.74 kg/h Oxygen =27884.01 kg/h Maleic anhydride =7158.16 kg/h Carbon monoxide =1168.42 kg/h Carbon dioxide = 1835.85 kg/h Water = 10021.49 kg/h Nitrogen =130843.36 kg/h
Maleic anhydride = 7086.58 kg/h Water = 100.21 kg/h Dibutyl phthalate = 20000 kg/h
12
xii
APPENDIX 8: MASS ENERGY BALANCE (STRIPPER)
Example Mass Balance Calculation for Deisobutanizer distillation column Calculation of component distribution using Hengstebeck Method
d log i m log ij C bi Calculation of saturation pressure of each component using the equation below:
The operating condition of the column is 128.30C and 0.066atm. Since, the pressure of the column is vacuum, a partial condenser is used to separate the vapor and liquid in the distillate. The light key (LK) is water distilled in the top at 99.9% and the heavy key (HK) is dibutyl phthalate distilled at the bottom at 99.9% The feed flow rate is 149.69 kgmole/h with the pre-specified feed composition at stream 12. Example calculation using maleic anhydride:
Partial pressure, Pi = Psat(x) = 9.88(0.48) = 4.715 kPa Volatility, Ki = Pi/PT = 4.715/6.666 = 0.71 Relative volatility, α = Ki/Khk = 0.71/0.016 = 290.02 The calculation of other component is done below: Component
Maleic
Mole
Mole
flow rate
fraction
Psat
Pi
Ki
α
Log α
72.27
0.482809
9.765642
4.71494
0.707299
290.0234
2.462
Water
5.562857
0.037163
255.2772
9.486962
1.423161
583.5581
2.766
DBP
71.85366
0.480028
0.033867
0.016257
0.002439
1
0
anhydride
A graph of log d/b and log α is plotted for the LK and HK as reference for finding the composition of the other components.
xiii
Example calculation for water (LK): d = 0.9999(5.56) = 5.56280 b = 0.001(5.56) = 0.00006 d/b = 5.56280/0.00006 = 99999 log d/b = 5 log α = 2.76608 Component
LK
HK
F
5.56286
71.85366
D
5.56280
0.00072
B
0.00006
71.85294
d/b
99999
0.00001
lg d/b
5
-5
lg α
2.76608
0
log d/b vs log α 6 y = 3.6152x - 5 4
log d/b
2 0 0.00000 -2
0.50000
1.00000
1.50000
2.00000
2.50000
-4 -6
log α
From the graph above, m = -3.6512 and c = -5. The d/b for other components are calculated. Example calculation for maleic anhydride:
Mol fraction going to bottom from feed, xbi = 1/(10672.46+1) = 0.0001 xiv
3.00000
Mol fraction going to distillate from feed, xdi = 1- xbi = 0.9999 Mol flow rate component in bottom, L = xbi(mol flow rate component) = 0.0001(72.27) = 0.013 kgmole/h Mol flow rate component in distillate, V = xdi(mol flow rate component) = 0.9999(72.27) = 72.2572 kgmole/h Component
log α
log di/bi
di/bi
xbi
xdi
V
L
Maleic
2.462
4.028
10672.464
0.000
1.000
72.263
0.007
Water
2.766
5.142
138538.604 0.000
1.000
5.563
0.000
DBP
0.000
-5.000
0.000
0.000
0.001
71.853
anhydride
1.000
Calculation for partial condenser at the distillate: Total moles entering flash drum at condenser = 74.69 kgmol/day Zj=mol fraction in feed, Xj=mol fraction in liquid in outlet, Yj=mol fraction in vapour in outlet,
V =vapour to feed ratio, Pj=vapour pressure of component supposed, P=Total pressure F
By hit and trail method At 85 oC and V/F=0.18 Components
Zj
Pj
Pj
KPa
P
Xj
Zj V 1 F
Pi P 1
Yi
Pj
Maleic anhydride
0.928
1.588
0.224
1.078
0.241
Water
0.072
57.767
8.136
0.032
0.257
Dibutyl phthalate
0.000
0.001
0.000
0.000
0.000
V=0.18*F =0.18(74.69) =13.444kgmol/day L=0.82*F = 0.82(74.69) =61.25 kgmol/day
xv
P
Xj
components
Liquid stream
vapour stream
kgmol
kgmol
L*Xj
V*Yj
Maleic anhydride
3.242
66.049
Water
3.462
1.938
Dibutyl phthalate
0.000
0.001
13 Maleic anhydride = 317.89 kg/h Water = 62.36 kg/h
14 Maleic anhydride = 7086.58 kg/h Water = 100.21 kg/h Dibutyl phthalate = 20000 kg/h
12
Maleic anhydride = 6476.54 kg/h Water = 34.92 kg/h Deisobutanizer column
15
xvi
Water = 0.000723366 kg/h Dibutyl phthalate = 20000 kg/h
APPENDIX 9: CP VALUE FOR ENERGY BALANCE CALCULATION
Compound
Propane
Isobutane
n-butane
Isobutene
1-butene
Neopentane
Iso-pentane n-pentane
Phase gas liquid solid gas liquid solid gas liquid solid gas liquid solid gas liquid solid gas liquid solid gas liquid solid gas
Tbp (0C)
Tc (K)
Pc (bar)
Hv [J/mol]
HF [J/mol]
-42.1
369.8
42.5
18786
-103920
-11.7
408.05
36.48
21399
-135600
-0.5
425.2
38
22408
-126230
-6.9
417.85
40.01
22100
-17900
-6.9
417.9
40
22131
-16910
9.5
433.78
31.99
22400
-167000
28
460.43
33.81
25220
-153700
36
469.6
33.7
25791
-146540
xvii
A 28.277 59.642 -11.23 6.772 71.791 110.211 20.056 62.873
Cp [joule/(mol K)] B C 1.16E-01 1.96E-04 3.28E-01 -1.54E-03 1.06E+00 -3.60E-03 3.41E-01 -1.03E-04 4.85E-01 -2.05E-03 -1.87E+00 1.44E-02 2.82E-01 -1.31E-05 5.89E-01 -2.36E-03
32.918 57.611 34.263 24.915 74.597 -11.985 -17.917 -186.315 105.567 -0.881 91.474 -10.547 26.671
1.85E-01 5.63E-01 6.85E-02 2.06E-01 3.34E-01 1.15E+00 5.72E-01 3.24E+00 -2.66E-01 4.75E-01 4.49E-01 1.25E+00 3.23E-01
7.79E-05 -2.30E-03 2.07E-03 5.98E-05 -1.39E-03 -3.58E-03 -4.17E-04 -1.09E-02 1.65E-03 -2.48E-04 -1.69E-03 -3.35E-03 4.28E-05
D -2.33E-07 3.65E-06 -3.68E-08 4.06E-06 -9.46E-08 4.23E-06 -1.46E-07 4.18E-06 -1.42E-07 3.02E-06 2.12E-07 1.34E-05 6.75E-08 3.13E-06 -1.66E-07
Oxygen
Maleic anhydride
Carbon monoxide
Carbon dioxide
Water
Nitrogen
liquid solid gas liquid solid gas liquid solid gas liquid solid gas liquid solid gas liquid solid gas liquid solid
-183
154.58
50.43
6820
0
200.2
721
72.8
54800
-470410
-191.37
132.92
34.99
6015.8
-110530
-56.57
304.19
73.82
25128
-394000
100
647.3
220.5
40683
-242000
-156
126.1
33.94
5570
0
xviii
80.641 -11.568 29.526 46.432 -16.683 -72.015 -12.662 32.5 29.556 125.595 21.83 27.437 -3981.02 41.195 33.933 92.053 9.695 29.342 76.452 24.334
6.22E-01 1.21E+00 -8.90E-03 3.95E-01 1.59E+00 1.04E+00 1.06E+00 2.10E-01 -6.58E-03 -1.70E+00 -4.71E-02 4.23E-02 5.25E+01 3.15E-02 -8.42E-03 -4.00E-02 7.50E-02 -3.54E-03 -3.52E-01 2.89E-01
-2.27E-03 -3.24E-03 3.81E-05 -7.05E-03 -2.99E-03 -1.87E-03 -2.32E-03 2.73E-04 2.01E-05 1.07E-02 6.41E-03 -1.96E-05 -2.27E-01 6.41E-05 2.99E-05 -2.11E-04 -1.56E-05 1.01E-05 -2.67E-03 1.16E-03
3.74E-06 -3.26E-08 3.99E-05 1.65E-06 2.05E-06 -1.22E-08 4.19E-06 4.00E-09 3.29E-04 -1.78E-08 5.35E-07 -4.31E-05 5.01E-05
Main Road
APPENDIX 10: DRAFT OF PLANT LAYOUT
Main Gate
Flare system area
Admin Building
Main Process Area
Room
Central Laboratory
l
Warehouse
Maintenance Building
Contro
Main sub statio n
Plant Gate
Expansion site
Waste water
Off site Utilities
treatment plant
Storage tank farm
xix
Main Road Plant Gate
Main Gate
Storage Tank Farm
Admin Building
Warehouse
Maintenance Building Central Laboratory Control Room
Expansion site Offsite Utilities Main Process Plant Waste water treatment plant
xx
Main sub station
