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final thesis

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final thesis

  1. 1. [Type text] FINAL YEAR PROJECT B.SC ENGINEERING PRODUCTION OF METHYL TERTIARY BUTYL ETHER USING SNAMPROGETTI PROCESS ADVISOR: MISS SANA ZAHID GROUP MEMBERS: FARAH NAEEM 2010-CH-06 ADEEBA KANWAL 2010-CH-10 IQRA JAMIL 2010-CH-14 ALI RAZA 2010-CH-82 AQSA ZULFIQAR 2010-CH-92 MARIA ASHRAF 2010-CH-116
  2. 2. 1 DEPARTMENT OF CHEMICAL ENGINEERING UNIVERSITY OF ENGINEERING AND TECHNOLGY LAHORE This project is submitted to the Department of Chemical Engineering, University of Engineering & Technology, Lahore, for the fulfillment of the requirements for the award of Bachelor’s degree in Chemical Engineering. Approved on: _______________________ Internal Examiner: _______________________ (Miss Sana Zahid) External Examiner: ________________________ Name: ________________________ DEPARTMENT OF CHEMICAL ENGINEERING UNIVERSITYOF ENGINEERING & TECHNOLOGY LAHORE PAKISTAN
  3. 3. 2 DEDICATION We dedicate this project to all the respected teachers of our department especially to our advisor Miss Sana Zahid who guided us a lot during whole project.
  4. 4. 3 Acknowledgment All praises to Almighty Allah, the most beneficent, the most merciful, Who guides us in the darkness and helps us in difficulties. Without His blessing, one cannot reach one’s destination. We are thankful to Allah Almighty who has blessed us with the courage, strength and wisdom so that we have been able to complete this project. In presenting this project report, we wish to express our profound feeling of gratitude to our respected supervisor, Miss Sana Zahid, Department Of Chemical Engineering, University of Engineering & Technology Lahore, without whose guidance and keen interest this work would not have been produced in the present form. We further take the opportunity to acknowledge the efforts and to express our deep appreciation to all the teaching staff of the department of chemical engineering who was every time ready to give us guidelines in spite of their busy schedule during the whole preparation of this report and provided us the necessary path for this completion. Here we must not forget to thank our parents who were always supportive and prayed for us every moment.
  5. 5. 4 Abstract The purpose for this MTBE or Methyl tertiary Butyl Ether plant is to produce 3,000,000 metric tonne/ year. MTBE is the simplest and most cost effective oxygenate to produce, transport and deliver to customers. The additive works by changing the oxygenate / fuel ratio so that gasoline burns cleaner, reducing exhaust emissions of carbon monoxide, hydrocarbons, oxides of nitrogen, fine particulates and toxic.. The raw materials used are isobutane, methanol, and water as feedstock. In addition, two types of catalysts are chromia alumina catalyzed compound in Snamprogetti Unit, while sulphonic ion exchanged resin catalyzed is used in the MTBE reactor. A good deal of catalyst has been devoted to improve the activity, selectivity, and the lifetime of the catalysts. In the Design Project, we emphasize in the individual designs for selected equipments in the plant. The chosen equipment are Catalytic Cracking Reactor, Multitubular Fixed Bed Reactor, MTBE Distillation Column, Liquid-Liquid Extraction Column and Heat Exchanger.This project also includes Process Control, Economic Evaluation, HAZOP study, EIA which are considered as group works.
  6. 6. 5 Contents 1 PROCESS BACKGROUND AND INTRODUCTION...................................................................................13 1.1 Introduction ................................................................................................................................13 1.2 Historical review of MTBE production process...........................................................................13 1.2.1 UOP-Oleflex Process ...........................................................................................................14 1.2.2 Philips Star Process .............................................................................................................14 1.2.3 ABB Lummus Catofin Process .............................................................................................14 1.2.4 Snamprogetti-Yartsingtez FBD (SP-Isoether)......................................................................15 2 PROCESS SELECTION ...........................................................................................................................15 2.1 Method Consideration................................................................................................................15 2.2 Detailed Process Description ......................................................................................................16 2.3 MTBE Unit ...................................................................................................................................16 2.4 Distillation Column Unit..............................................................................................................17 2.5 Liquid-Liquid Extraction Unit.......................................................................................................17 3 MATERIAL AND ENERGY BALANCE......................................................................................................18 3.1 Snamprogetti Unit (Reactor and Regenerator)...........................................................................18 3.2 SEPARATOR.................................................................................................................................21 3.3 MIXING POINT.............................................................................................................................22 3.4 MTBE Reactor..............................................................................................................................24 3.5 Material Balance across 1st Distillation Column:.........................................................................26 3.6 Liquid-Liquid Extraction Column.................................................................................................27 3.7 Distillation Column-II ..................................................................................................................28 4 DESIGN PROJECT.................................................................................................................................30 4.1 Catalytic Cracking Design............................................................................................................30 4.1.1 Introduction ........................................................................................................................30 4.1.2 Estimation of Diameter of Reactor .....................................................................................31 4.1.3 Calculation of the Transport disengagement height, TDH .................................................32 4.1.4 Minimum Fluidization Velocity ...........................................................................................32 4.1.5 Calculation For The Value Of Terminal Velocity Ut .............................................................32 4.1.6 Solids Loading......................................................................................................................33 4.1.7 Calculation of Residence Time............................................................................................34
  7. 7. 6 4.1.8 Calculation For The Pressure Drop......................................................................................34 4.1.9 Specification Sheet:.............................................................................................................35 4.2 DESIGN OF FIXED BED MULTITUBULAR REACTOR......................................................................35 4.2.1 Reaction Taking Place on the Catalyst ................................................................................35 4.2.2 Catalyst Properties:.............................................................................................................36 4.2.3 Steps for design:..................................................................................................................36 4.2.4 Particle solid density:..........................................................................................................36 4.2.5 Void Volume of catalyst:.....................................................................................................36 4.2.6 Reaction Rate:.....................................................................................................................37 4.2.7 Weight of catalyst: ..............................................................................................................37 4.2.8 Volume of the catalyst bed:................................................................................................37 4.2.9 Pressure drop in the bed:....................................................................................................37 4.2.10 Height of the bed: ...............................................................................................................37 4.2.11 Total cross section area of tubes: .......................................................................................37 4.2.12 Tube diameter:....................................................................................................................38 4.2.13 Tube area: ...........................................................................................................................38 4.2.14 Number of tubes:................................................................................................................38 4.2.15 Pitch of the tube: ................................................................................................................38 4.2.16 Bundle diameter: ................................................................................................................38 4.2.17 Shell Diameter:....................................................................................................................38 4.2.18 Shell side heat transfer co efficient: ...................................................................................38 4.2.19 Tube side co efficient:.........................................................................................................40 4.2.20 Overall Heat transfer Co efficient: ......................................................................................41 4.2.21 Specification Sheet:.............................................................................................................41 4.3 DESIGN OF DISTILLATION COLUMN............................................................................................42 4.3.1 CHOICE BETWEEN PLATE AND PACKED COLUMN ..............................................................42 4.3.2 Choice of Plate Type............................................................................................................43 4.3.3 DESIGNING STEPS OF DISTILLATION COLUMN ...................................................................44 4.3.4 Design Data.........................................................................................................................45 4.3.5 Rmin Using Underwood Method...........................................................................................47 4.3.6 Calculating Minimum No. Of Stages ...................................................................................47 4.3.7 Theoretical Number Of Stages............................................................................................48 4.3.8 Actual Number Of Stages:...................................................................................................48 4.3.9 Determination Of The Column Diameter............................................................................48
  8. 8. 7 4.3.10 Vapour Load At Bottom ......................................................................................................48 4.3.11 Liquid Load at Bottom.........................................................................................................49 4.3.12 Tray Dynamics.....................................................................................................................49 4.4 Liquid-liquid extraction unit design ............................................................................................52 4.4.1 Choice Of Solvent................................................................................................................52 4.4.2 Calculation of number of stages .........................................................................................53 4.4.3 Sizing of the sieve tray ........................................................................................................54 4.5 Design of a distillation column:...................................................................................................59 4.5.1 Calculation of Minimum no. of Plates.................................................................................59 4.5.2 Calculation of Minimum Reflux Ratio .................................................................................60 4.5.3 Actual Reflux Ratio:.............................................................................................................60 4.5.4 Theoretical no. of Plates: ....................................................................................................60 4.5.5 Actual number of stages:....................................................................................................61 4.5.6 Number of plates in rectification and stripping section: ....................................................61 4.5.7 Diameter of column ............................................................................................................61 4.5.8 Weir calculations.................................................................................................................62 4.5.9 Pressure drop......................................................................................................................63 4.5.10 Height of column.................................................................................................................64 4.5.11 Specification sheet of distillation column:..........................................................................66 4.6 DESIGN OF HEAT EXCHANGER: ...................................................................................................67 4.6.1 INTRODUCTION...................................................................................................................67 4.6.2 Calculation of Heat Duty.....................................................................................................70 4.6.3 Calculation of LMTD............................................................................................................70 4.6.4 Overall Heat Transfer Coefficient (U)..................................................................................70 4.6.5 Calculation of heat transfer Area........................................................................................70 4.6.6 Number of Tubes Calculation .............................................................................................70 4.6.7 Number of Tubes ................................................................................................................71 4.6.8 Bundle and Shell Diameter Calculation...............................................................................71 4.6.9 Tube Side Coefficient , hi.....................................................................................................72 4.6.10 Reynolds And Prandtel Number..........................................................................................72 4.6.11 Shell Side Coefficient, hs......................................................................................................73 4.6.12 Shell Side Equivalent Diameter for Triangular Pitch Arrangement.....................................73 4.6.13 Shell Side Coefficient, hs......................................................................................................74 4.6.14 Overall Heat Transfer Coefficient, Uo.................................................................................74
  9. 9. 8 4.6.15 Tube Side Pressure Drop.....................................................................................................75 4.6.16 Shell Side Pressure Drop .....................................................................................................75 4.6.17 Specification Sheet..............................................................................................................75 5 HYSYS ..................................................................................................................................................76 5.1 The Design Based On Hysys Simulation......................................................................................76 5.2 Results.........................................................................................................................................79 6 ENVIRONMENTAL IMPACT ASSESSMENT ...........................................................................................80 6.1 INTRODUCTION...........................................................................................................................80 6.2 Steps of EIA: ................................................................................................................................82 6.3 Impacts of project on environment:...........................................................................................82 6.3.1 STACK GASES.......................................................................................................................82 6.3.2 WASTEWATER TREATMENT................................................................................................83 6.3.3 Waste Minimization............................................................................................................86 7 HAZOP STUDY .....................................................................................................................................87 7.1 Introduction ................................................................................................................................87 7.2 Steps Conducted In Hazop Study................................................................................................87 7.3 HAZOP METHOD FLOW DIAGRAM..............................................................................................88 7.4 HAZOP Guide Words and Meanings ...........................................................................................89 8 Economic Analysis...............................................................................................................................93 8.1 Global Market .............................................................................................................................93 8.2 Asia Market.................................................................................................................................93 8.3 Worldwide Demand....................................................................................................................94 8.4 Market Price................................................................................................................................94 8.4.1 Methanol.............................................................................................................................94 8.4.2 Isobutane ............................................................................................................................95 8.5 Cost estimation...........................................................................................................................95 8.5.1 Direct Cost...........................................................................................................................95 8.5.2 Indirect Cost........................................................................................................................96 8.5.3 Total Capital Investment.....................................................................................................96 8.5.4 Production Cost...................................................................................................................96 8.5.5 Manufacturing Cost ............................................................................................................97 8.5.6 General Expenses................................................................................................................97 8.5.7 Distribution and Selling Cost...............................................................................................97 8.5.8 Research and Development Cost........................................................................................97
  10. 10. 9 8.5.9 Financing (Interest).............................................................................................................98 9 INSTRUMENTATION AND CONTROL ...................................................................................................98 9.1 Introduction ................................................................................................................................98 9.2 OBJECTIVES OF CONTROL ...........................................................................................................98 10 References ....................................................................................................................................101
  11. 11. 10 NOMEN CLATURE Ar - Archimedes number a - acceleration B - settling chamber longitudinal cross-sectional area b - dimension C - constant CD - drag coefficient c - concentration D - system diameter d - particle diameter de - effective fiber diameter E, - field intensity F - cross-sectional area Pr - Fronde number g - gravitational acceleration H - height K - precipitation constant , A - Cross sectional area of catalytic reactor Aor - Area of orifice Bv d - Diameter of bubble in the bed dp - Particle diameter D - Diffusivity Dt - Diameter of catalytic reactor e - Thickness E - Activation energy FBo - Mass flow of coal to the catalytic reactor FC - Fixed carbon mass fraction Hbed - Height of bed Hh - Height of Catalytic reactor J - Joint factor
  12. 12. 11 k” - Reaction rate constant k - Reaction rate constant eq K - Equilibrium constant L - Height above the bed n - Total no of orifice N - No of holes in 1 m2 area Nor - No of orifice in 1 m2 area Pi - Design stress rC , rS - Rate of reaction R - Ideal gas constant Ret - Reynolds number Rp - Radius of particle t - Total holding time T - Temperature Uo - Superficial gas velocity Umf - Minimum fluidization velocity Ut - Terminal velocity VBed - Volume of bed WBed - Weight of coal in bed WC - Total mass of carbon X - Conversion factor a - Fitting parameter (for this design is 0.21) b - Fitting parameter (for this design is 0.66) rg - Gas density rB - Molar density rs - Bulk density of catalyst rp - Particle density μg - Gas viscosity t - Time for complete conversion of reactant particle Dp - Pressure drop E - total elutriation rate of particles Ef - frictional force of particles
  13. 13. 12 Ei - entrainment rate of panicle size i Ei∞ - elutriation rate of particle size i Eo - total entrainment rate at bed surface E∞ - total elutriation rate of particles g - gravitational acceleration constant gc - gravitational conversion constant, m kg/s2 kg -force Gi - solids flow rate h - height above dense bed surface Rep - particle Reynolds Number = r ( ) /m g o ts p U -U d Ret - dpUrg /m t - time Umf - minimum fluidization velocity Uo - superficial gas velocity Usi - solid velocity (upward) Us - single particle terminal velocity of particle size i W - weight fraction of bed Ws - weight of solid particles in verlical pipe having length h Xi - weight fraction of particle size i in bed Greek Symbols e - voidage in freeboard - voidage in freeboard for system having only particle size i l - solid friciion coefficient g r- gas density p r- particle density
  14. 14. 13 1 PROCESS BACKGROUND AND INTRODUCTION 1.1 Introduction Methyl tertiary butyl ether (MTBE) is produced by reacting isobutene with methanol over a catalyst bed in the liquid phase under mild temperature and pressure. Isobutene can be obtained from stream cracker raffinate or by the dehydrogenation of isobutane from refineries. Ether in general is a compound containing an oxygen atom bonded to two carbon atoms. In MTBE one carbon atom is that of a methyl group – CH3 and the other is the central atom of a tertiary butyl group, -C (CH3). At room temperature, MTBE is a volatile, flammable, colorless liquid with a distinctive odor. It is miscible with water but at high concentrations it will form an air-vapor explosive mixture above the water, which can ignite by sparks or contact with hot surfaces. MTBE has good blending properties and about 95% of its output is used in gasoline as an octane booster and an oxygenate (providing oxygen for cleaner combustion and reduced carbon monoxide emissions). It is also used to produce pure isobutene from C4 streams by reversing its formation reaction. It is a good solvent and also can be used for extraction. MTBE properties 88.15Molecular weight (g/mole) 54Boiling Temperature ( °C) 0.74Specific gravity 50000Water solubility (mg/L) 251Vapor pressure (mm Hg) 1.5E-3Henry's Law 81.7Latent heat of vaporization (Cal/g) 0.51Specific heat (Cal/g.°C) 1.2 Historical review of MTBE production process The MTBE plants actually consist of six units: Isomerization Unit (including deisobutanizer), Dehydrogenation Unit, MTBE Unit, Methanol Recovery Unit, Oxygenate Removal Unit and Olefin Saturation Unit. A common offsite utility system will be incorporated to distribute the required utilities to each unit. There are four method of producing MTBE implemented under license as the following:
  15. 15. 14 1. UOP-Oleflex Process 2. Phillips STAR Process 3. ABB Lummus Catofin Process 4. Snamprogetti-Yarsingtez FBD (SP-Isoether) Process. 1.2.1 UOP-Oleflex Process The UOP-Oleflex process uses multiple side-by-side, radial flow, moving-bed reactors connected in series. Preheated feed and interstage heaters supply the heat of reaction. The reaction is carried out over platinum supported on alumina, under near isothermal conditions. The catalyst system employs UOP's Continuous Catalyst Regeneration (CCR) technology. The bed of catalyst slowly flows concurrently with the reactants and is removed from the last reactor and regenerated in a separate section. The reconditioned catalyst is then returned to the top of the first reactor. The typical processes involved are the deisobutenization, the isomerisation and the dehydrogenation process that has been commercial in Malaysia 1.2.2 Philips Star Process The second one is the Philips Steam Active Reforming (STAR) Process. The Phillips Steam Active Reforming (STAR) Process uses a noble metal-promoted zinc aluminate spinel catalyst in a fixed-bed reactor. The reaction is carried out with steam in tubes that are packed with catalyst and located in a furnace. The catalyst is a solid, particulate noble metal. Steam is added to the hydrocarbon feed to provide heat to the endothermic reaction, to suppress coke formation, and to increase the equilibrium conversion by lowering partial pressures of hydrogen and propane. 1.2.3 ABB Lummus Catofin Process The ABB Lummus Catofin Process uses a relatively inexpensive and durable chromium oxide–alumina as catalyst. This catalyst can be easily and rapidly regenerated under severe conditions without loss in activity. Dehydrogenation is carried out in the gas phase over fixed beds. Because the catalyst cokes up rapidly, five reactors are typically used. Two are on stream, while two are being regenerated and one is being purged. The reactors are cycled between the reaction and the reheat/regeneration modes, and the thermal inertia of the catalyst controls the cycle time, which is typically less than 10 minutes. The chromium catalyst is reduced from Cr6+ to Cr3+ during the dehydrogenation cycle. The raw materials used to produce MTBE by using this method are butanes, hydrogen and as well as recycled isobutene from the system itself. In this process, there is an isostripper column, which separates the heavies, and the light ends from which then could produce MTBE.
  16. 16. 15 1.2.4 Snamprogetti-Yartsingtez FBD (SP-Isoether) The Snamprogetti-Yarsingtez SP-Isoether (FBD) Process uses a chromium catalyst in equipment, which is the fluidized bed that resembles conventional fluidized catalytic cracking technology used in the oil refinery. The catalyst is recirculated from the reactor to the regeneration section on a 30–60-min cycle. The process operates under low pressure and has a low-pressure drop and uniform temperature profile. Snamprogetti has been presenting and marketing their hydrogenation technology, ISOETHER 100, since 1997. This process is to be used to convert MTBE units by utilizing Snamprogetti’s MTBE Water Cooled Tubular Reactor Technology. In this SP-Isoether Process, the products are MTBE and isooctagenas (iso octane gas). In this SP-Isoether Process the catalyst used in the isoetherification reactor is the same as those other typical processes, which is Platinum. 2 PROCESS SELECTION Suitable process, which is gives a lot of profit and less problem is an important in order to determinant for the success of a plant. This chapter will briefly discuss the best process selected based on a few criteria. It covers general consideration, detailed consideration for process selection and conclusion on the process selection. 2.1 Method Consideration From the processes mentioned earlier, there are many ways to produce MTBE. It is essential to choose the best method that will be used to produce MTBE. The selection of the method must consider the safety of the plant, minimum waste or by product generated, efficient an economical. Snamprogetti-Yarsingtez SP- Isoether FBD process will be chosen as the method to produce MTBE. More detailed reasons for the selection of this process are: High conversion (greater than 98 %) with few by-products compared to other process. From the economy aspect,Snamprogetti-Yarsingtez FBD Process can reduce the cost of setting up the plant as it can be implied in any of typical MTBE-produced plant, known as “Financial Safety Net”.(When an MTBE plant faces an oversupplied MTBE market, Isoether makes it possible to switch production from MTBE to a superior Alkylate.). As for the safety aspects of the plant, as the Snamprogetti-Yarsingtez FBD is a safe process as it just use the fluidize bed to the process of producing MTBE. The process operates under low pressure and has a low pressure drop and this means that the fluidized bed is physically not harmful to anyone. As for the temperature, it operates under uniform temperature profile. As the temperature is not high, this means that the process is not as dangerous as other high temperature-operated process. But, precautions should be taken seriously all the time, as we do not know when an accident could happen even in the safest place. As for the waste by using the Snamprogetti-Yarsingtez FBD Process, the product of the process is only MTBE and other effluent and as well as flue gas which are not harmful to the environment.
  17. 17. 16 2.2 Detailed Process Description The Snamprogetti-Yarsingtez SP-Isoether (FBD) Process uses a chromium catalyst in equipment, which is the fluidized bed that resembles conventional fluidized catalytic cracking technology used in the oil refinery with 65% isobutane (i-C4H10) conversion to produce isobutene. Dehydrogenation reaction that occur in this process: iC4H10 iC4H8 + H2 The main feature of this process is that the catalyst filled annuli are connected in such a way that small, discrete amounts of catalyst can be withdrawn from the bottom of a reactor, and sent to the top of the reactor. Catalyst withdrawn from the bottom of the reactor is sent to a separate regeneration section for regeneration prior to being sent to the top of the reactor. The catalyst is recirculated from the reactor to the regeneration section on a 30–60-min cycle. The reactor and regeneration sections are totally independent of each other. The regeneration section can be stopped, even for several days, without interrupting the dehydrogenation process in the reactor section. The vaporized isobutane is fed along with fresh catalyst to the first, called reactor, and the spent catalyst is separated from the products and sent to the regenerator, where air (O2) is added to oxidize the carbon. The reactor cracks the isobutane and forms coke on the catalyst. Then in the regenerator the coke is burned off and the catalyst is sent back into the reactor. The “magic” of this process is that the reactor-regenerator combination solves both the heat management and coking problems simultaneously. Burning off the coke is strongly exothermic, and this reaction in the regenerator supplies the heat (carried with the hot regenerated catalyst particles) for the endothermic cracking reactions in the reactor. The process operates under low pressure and has a low-pressure drop and uniform temperature profile. Products that have been produced from this unit are isobutene. Isobutene available in the C4 stream from the Snamprogetti-Yarsintez FBD unit will be combining with methanol, which is sourced from the Sabah Gas Industries methanol plant in Labuan to produce, fuel-grade MTBE with a high-octane value in the MTBE unit. 2.3 MTBE Unit The MTBE unit includes two sections such as the main reaction section and the finishing reaction. In the main reaction section, 98% conversions of isobutene occurs mainly in the main reactor which are designed to provide the mechanical ands thermal conditions required by the expanded catalyst bid technology. Reactions occur in this unit are 1. iC4H8 + CH3OH C5H12O ( isobutene) (methanol) (MTBE)
  18. 18. 17 2. CH3OH + CH3OH (CH3)2O + H2O (DME) 3. iC4H8 + H2O C4H10O (TBA) The reactor is operated in an up-flow direction with an external liquid recycle to remove the heat of reaction and to control the expansion of the catalyst bed. This selective reaction of methanol with isobutene is conducted in liquid phase at moderate temperature on an ion exchange resin type catalyst. The expansion of the catalyst bed in the reactor is ensured by pump around circulation loop with a cooling water cooler to control the reactor feed temperature to remove the heat of reaction. Resin traps on top of each reactor to trap resin in case of carryover with the liquid. In the finishing reactor section, isobutene final conversion is achieved in a catalytic column where reaction and distillation are performed simultaneously. 2.4 Distillation Column Unit This column includes a separation column yielding MTBE product at the bottom and (isobutene, isobutene, normal butane, water and DME) with methanol entrained by azeotropy at the top. The reaction section bed is contained in the upper part of this column. An excess of methanol is maintained corresponding to the amount leaving the tower in the azeotrope. The required methanol is passed through guard beds and filtered prior to being charged to the catalytic column to achieve final conversion. Bottom MTBE product and the other by-product such as TBA, DME is sent to rundown tanks under level control after cooling in feed/bottom exchanger and trim cooler. The overhead of the column is condensed in the air-cooled condenser under pressure control. One part of the liquid is sent to the column as reflux and the other part to the liquid- liquid extraction unit after cooling. 2.5 Liquid-Liquid Extraction Unit In this unit methanol will extract from the isobutene, isobutene, normal butane to produce C4 raffinate from the overhead of the column and at the bottom, methanol and water are produced. C4 raffinate from this unit we decided to sell to the Korea.
  19. 19. 18 3 MATERIAL AND ENERGY BALANCE Basis of Material & Energy Balance  Steady state  Capacity of MTBE plant  7920 hours/yr  3000000 Tones/yr  Reference temperature  25⁰C  Reference pressure  101.325 Kpa  Conversion in 1st Reactor 65%  Overall Conversion of isobutene in 2nd reactor 98%  Yield of MTBE 0.95  Conversion of methanol into TBA 98% 3.1 Snamprogetti Unit (Reactor and Regenerator) The fraction at stream S3 acquired from isobutane instrument grade, MSDS. Reaction occurred in the reactor, iC4H10 C4H8 + H2
  20. 20. 19 Flowrate in kgmole/hr of iC4H10 in the feed stream of S3 = 5.56x107 Balanced Based upon the stoichoimetric ratio with 65% conversion of iC4H10 to obtain C4H8. Since, 65% conversion in the reactor, kgmole/hr of C4H8 obtained = 0.65*5.56x107 =3.58x107 kgmole/hr 35% of iC4H10 unreacted = 1.928x107 kgmole/hr Based upon stoichiometric ratio (inert) (unreacted) (inert) n C4H10 + iC4H10 C4H8 + H2 + iC4H10 + n C4H10 Inputs S3 Output S4 Components Mass Flow kg/yr Mass Flow kg/yr i-butane 1.3x107 1.1x109 n-butane 3.196x109 1.3x107 H2 - 7.2x107 i-butene - 2.01x109 Total 3.2x109 3.2x109 For energy balance of the reactor Cp values can be calculated using the formula CP = a + bT + cT2 + dT3 For this sample of calculations, listed are the values of constants (a, b, c, d ) in the ideal gas heat capacity equation based on R. K Sinnot, Coulson & Richardson, Chemical Engineering, Volume 6, Third Edition, Butterworth Heinemann: Components a KJ/Kmol B KJ/Kmol C KJ/Kmol d KJ/Kmol ΔĤF KJ/Kmol C5H12O 2.53 5.14 x 10-1 -2.60 x 10-4 4.30 x 10-8 -292990
  21. 21. 20 CH3OH 2.12 x 101 7.09 x 10-2 2.59 x 10-5 -2.85 x 10-8 -201300 H2O 32.243 1.93 x 10-3 1.06 x 10-5 - -3.60 x 10-9 -242000 C4H8 -2.994 3.53 x 10-1 -1.98 x 10-4 4.46 x 10-8 -130 i-C4H8 16.052 2.8043 x 10-1 -1.091 x 10-4 i9.098 x 10-9 i-16900 i-C4H10 -1.39 3.85 x 10-1 -1.85 x 10-4 2.90 x 10-8 --134610 C4H10O -4.86 x 101 7.17 x 10-1 --7.08 x 10-4 2.92 x 10-7 --312630 n-C4H10 9.85 - 3.31 x 10-1 - 1.11 x 10-4 -2.82 x 10-9 -126.23 (CH3)2O 1.70 x 101 1.79 x 10-1 -5.23 x 10-5 -1.92 x 10-9 -184180 H2 2.71 x 101 9.27 x 10-3 -1.38 x 10-5 7.65 x 10-9 0 ΔĤR = (ΔĤF C4H8) + (ΔĤF H2) -(ΔĤF i-C4H10) ΔH in = mCp ΔT ΔH out = mCp ΔT Q = ΔHR + (ΔHout) - (ΔHin) = 3.66x1016 KJ/yr ΔH in ΔHr ΔH out Enthalpy KJ/yr 7.83x1013 3.34x1015 1.10x1014
  22. 22. 21 3.2 SEPARATOR OutputInput S11S11S10S10S9S9Stream Mass flow (Kg/Yr) Molar flow (Kgmole/Yr) Mass flow (Kg/Yr) Molar flow (Kgmole/Yr) Mass flow (Kg/Yr) Molar flow (kgmole/Yr) Component 2.01×109 35819153--2.01×109 35819153Butene --71638306358191537163830635819153Hydrogen 1.12×109 19287236--1.12×109 19287236i-Butane 12836026221310.8--12836026221310.8n-Butane 3.14×109 71683063.21×109 Total Energy balance around separator Stream 9 H (KJ/hr)∆T (K)To (K)Hf(KJ/Kmol)∆Components 166.03326.3298-1.69×104 Butene
  23. 23. 22 2.17326.32981.35×105 i-Butane 353.50326.3298-1.26×105 n-Butane 99.88326.32980.00Hydrogen Stream 10 Stream 11 H (KJ/hr)∆T (K)To (K)Hf(KJ/Kmol)∆Components 321.92326.3298-16910Butene 188.89326.3298-134610i-Butane 2.18326.3298-126230n-Butane Q = (∑H)out – (∑H)in = -208.46 kW 3.3 MIXING POINT OutputInput S14S27S13Stream Mass flowMolar flowMass flowMolar flowMassMolar FlowComponents H (KJ/hr)∆T (K)To (K)Hf(KJ/Kmol)∆Components 99.8326.32980.00Hydrogen
  24. 24. 23 (Kg/yr)(Kgmole/y r) (Kg/Yr)(Kgmole/yr)flow (Kg/yr) (Kgmole/yr) 126083418539401068259197980999.341.26×109 Methanol 0.072158237.22596484144249.1--Water 126083418551884631.26×109 Total Energy balance around mixing point Stream 13 H (KJ/hr)∆T (K)To (K)Hf(KJ/Kmol)∆Component 11.79300298-2.01×102 Methanol Stream 27 H (KJ/hr)∆T (K)To (K)Hf(KJ/Kmol)∆Components 0.02300298-2.01×105 Methanol 0.03300298-2.42×105 Water Stream 14 H (KJ/hr)∆T (K)To (K)Hf(KJ/Kmol)∆Components 11.81300298-2.01×105 Methanol 0.03300298-2.42×105 Water Q = (∑H)out – (∑H)in = 0.04 kW
  25. 25. 24 3.4 MTBE Reactor Material Balance There are three reactions taking place in MTBE reactor: I. i-C4H8+CH3OH→C5H12O II. 2CH3OH→C2H6O+H2O III. i-C4H8+H20→C4H10O Reaction 1 is main reaction and 2nd and 3rd reactions are side reactions.  The conversion of i-butene is 98%.  The yield of MTBE is 0.95  The conversion of methanol into DME is 98%. COMPONENT MASS IN (Kg/YEAR) MASS OUT (Kg/YEAR) i-butene 2.01x1009 40117451 n-butane 12836026 12836026
  26. 26. 25 i-butane 1.12x1009 1.12x1009 Methanol 1.26x1009 3438639 Water 2596484 28046397 MTBE 0 2.99x1009 TBA 0 79518520 DME 0 1.21x1008 Total 4.4x1009 4.4x1009 Energy balance PARAMETER S11 S14 S15 TEMPERATURE (⁰C) 53.3 27 101 PRESSURE(kPa) 2000 2000 1950 By using the values of cp and formula for calculating heat of reaction: ∆HR=∆Hf for products -∆Hf for reactants We get: • ∆HR1= -91999.877 kJ/mol • ∆HR2= -26000 kJ/mol • ∆HR3= -68999.877 kJ/mol Total Q=-2.1x1010 kJ/year
  27. 27. 26 3.5 Material Balance across 1st Distillation Column: Methanol recovery in top= 0.93 Water recovery in Bottom= 0.65 Components Mass in(kg/year) Top mass(kg/year) Bottom mass (kg/year) i-butene 4.01x107 4.01x107 0 n-butane 1.28x107 1.28x107 0 i-butane 1.12x109 1.12x109 0 Methanol 3.43x106 3.20x106 2.41x105 MTBE 2.99x109 0 2.99x109 Water 2.8x107 9.82x106 3.24x107 DME 1.21x108 1.21x108 0 TBA 7.95x107 0 7.95x107 Total 4.4x109 1.31x109 3.11x109
  28. 28. 27 ENERGY BALANCE: PARAMETER FEED TOP BOTTOM TEMPERATURE (⁰C) 64.5 22 121.7 PRESSURE (kPa) 500 355 570 Enthalpies Feed= -2.13x1009 kJ/year Top stream= -1.25x1009 kJ/year Bottom stream= -2.39x1009 kJ/year 3.6 Liquid-Liquid Extraction Column Stream 6 consists of water only. We are adding water because methanol is soluble in water. Components Stream 1 (kg/yr) Stream 2 (kg/yr) Stream 3 (kg/yr) Stream 4 (kg/yr) CH3OH - 3.20x1006 3.20x1005 2.88x1006
  29. 29. 28 H2O 1.20x10 05 9.82x1006 - 1.30x1008 n-C4H10 - 1.28x1007 1.28x1007 - i-C4H10 - 1.12x1009 1.12x1009 - i-C4H8 - 4.01x1007 4.01x1007 - DME - 1.21x1008 1.21x1008 - Total Flow rate (kg/yr) 1.20x1005 1.31x1009 1.29x1009 1.33x10 08 Temperature(⁰C) 27 22 25 141 Pressure( kPa) 100 355 377 377 ∆H (kJ/yr) - 1.512x1016 -2.890x1016 -2.827x1016 -1.48x1016 3.7 Distillation Column-II
  30. 30. 29 Here the top product is the recycled to the mixer, basically we are recovering methanol which is raw material.
  31. 31. 30 4 DESIGN PROJECT 4.1 Catalytic Cracking Design 4.1.1 Introduction A bed of solid particles can be fluidized by a stream of gas through it. The fluidization of solids in a stream of gas occurs only if the gas velocity achieved a certain value which is called minimum fluidization velocity Umf. Once the gas velocity achieved this value, the bed expands and pressure drop across the fluidized bed remains constant once fluidization occurred. In this commercial fluidized-bed catalytic cracking reactor, catalysts flow up through the reaction regeneration section in a riser type of flow regime. The over head catalyst captured by cyclones is returned to the hopper where it is fluidized with air to recapture any entrained hydrocarbon vapor. The catalyst was then discharged from the hopper,down through a standpipe. The solids flow through the standpipe was controlled by slide valve located at the base. From there, the solids went into the riser where they are carried by stream of air to the regenerator vessel. The regenerator operation in these plants resembled that of the reactor except for the system’s use of air instead of oil vapor. A portion of the catalyst from the regenerated catalyst hopper was returned to the regenerator through catalyst fresh feed exchangers. This action controlled the regenerator temperature and served to preheat the feed. Another bypass line from Components Stream 1(kg/yr) Stream 2 (kg/yr) Stream 3 (kg/yr) CH3OH 2.88x1006 2.59x1006 2.86x1005 H2O 1.30x1008 2.60x1006 1.27x1008 Total flowrate (kg/yr) 1.33x1008 5.18x1006 1.27x1008 Temperature (⁰C) 141 126 155 Pressure (kPa) 377 246 543 ∆H kJ/yr) -2.19x1009 -2.09x1009 -1.84x1009
  32. 32. 31 the hopper to the regenerator was used to control the dense bed level or holdup in the regenerator. Catalyst from the regenerated catalyst hopper flowed through a standpipe back into riser where the feed was injected. The commercial cracking catalysts used most widely is silica- alumina. High content catalysts are characterized by higher equilibrium activity level and surface area. These catalysts could be offered at a lower price. An advantage of this catalyst grade is that a lesser amount of adsorbed, unconverted, heavy products on the catalyst were carried over to the stripper zone and regenerator. As a result, a higher yield of more valuable products and also smoother operation of the regenerator was achieved. Basically the design of the fluidized bed system can be divided into several sections:  Estimation of diameter of reactor  Calculation of the Transport disengagement height, TDH  Minimum fluidization Velocity  Calculation for the value of terminal velocity Ut  Calculation of solid loading  Calculation of residence time  Calculation of bed height  Calculation for the pressure drop ∆PB Available data Dp = diameter of particle = 8010-6 m ρp = density of particle = 1282 kg/ m3 ρg = density of gas = 1.484 kg/ m3 μ = 1.15x10-5 Flow rate of gas stream = 39353 kg/hr 4.1.2 Estimation of Diameter of Reactor The fluidized bed diameter depends on the operating gas velocity. A larger diameter is required for a low gas velocity while for a high gas velocity, a small diameter is required. However the gas velocity must exceed the terminal velocity (Ut) of the particle transport of solid particles may occur. The operating velocity should be between minimum fluidization velocity and terminal falling velocity to maintain fluidization of solids. 𝑈 𝑜 = 𝐷 𝑃 2 g(𝜌 𝑝 − 𝜌 𝑔) 18μ Uo = 0.388 m/s
  33. 33. 32 The bed diameter will be depending on the area of reactor used: Cross sectional area, A = Q/ Uo A = 18.985 m2 diameter of bed = √ 4𝐴 𝜋 3.29m 4.1.3 Calculation of the Transport disengagement height, TDH According to M. Rhodes (1998), the TDH region is considered as the region where located above the bed surface to the top of disengagement zone. While the disengagement zone is the region above the splash zone or region just above the bed surface in which the upward flux and suspension concentration of fine particles decrease with increasing in height. There are so many correlations that can be used to find the TDH value. For this design Amitin et al. (1968) was used. TDH (F) = 0.85Uo 1.2 (7.33 -1.2log10Uo) = 8m 4.1.4 Minimum Fluidization Velocity According to the Martin Rhodes (1999) Re = 𝑈 𝑚𝑓𝐷 𝑝𝜌 𝑔 𝜇 = 𝜇 𝜌𝑑𝑝 (1135 .7 + 0.0408 Ar )0.5 -33.7 Umf = 0.425 m/s 4.1.5 Calculation For The Value Of Terminal Velocity Ut From the chart of CDRet2 and CD/Ret VS Reynolds number, for value of
  34. 34. 33 𝐶 𝐷Re2 = 4𝜌 𝑔(𝜌 𝑝 − 𝜌 𝑔)gdp3 3μ2 Ut = 0.20 m/s 4.1.6 Solids Loading Solids loading unreturned = mass flow of solids volume flow of gas = 𝑅 𝑈𝐴 = 0.081kg/s R is calculated by through iteration, 𝑚 𝑏𝑖 = 𝐹 𝑚𝑓𝑖 (F−R)+𝐾 𝐼𝐴
  35. 35. 34 Solids loading return = F-R = 8.915 kg/s 4.1.7 Calculation of Residence Time The outlet concentration of a plug flow reactor is related to the inlet concentration of the reactant by the same equation as in a batch reactor with the same residence time.For an equilibrium reaction between A and B, is first order. Based on studied of Khabtou, S., Chevreau, T., and Lavalley, J.C., Micropor. Mat. 3,133 (1994), express the rate per catalyst mass instead of reactor volume. 𝐾 𝑚𝑎 = − 1 ST K K + 1 ln [1 − x − x K ] The value K = 0.75 and ST = 151.94 g.hr/m3 based on study by Yamamoto, S. Asaoka, et al (1997). Kτ = (1 − ɛ)ln 1 1 − 𝑋 − ɛx 𝛕 = 115.99s t = Wbed FBo Wbed = 9567.91 kg Volume of the bed = weight of bed particle bulk density = 74.6 m3 Height of bed = Weight of bed (1−ɛ)Aρ = 6.6m 4.1.8 Calculation For The Pressure Drop The equation that can be used to calculate the pressure drop across the bed is: ∆P = 𝑊 𝑏𝑒𝑑−𝑊 𝑏𝑒𝑑( r𝑔 r𝑝 ) 𝐴 ∆P = 49.383 Kpa
  36. 36. 35 4.1.9 Specification Sheet: Components Estimated values Min. fluidization Velocity 0.00425m/s Catalyst type Silica-alumina Catalyst weight 95679.771 kg Diameter of catalyst bed ( reactor) 3.29 m Height of reactor 6.6m Volume of reactor 74.6 m3 Pressure drop 49.38 KPa Materials of construction Stainless Steel (18Cr/8Ni, 304) 4.2 DESIGN OF FIXED BED MULTITUBULAR REACTOR Fixed bed reactors are the most important type of the reactor for the synthesis of large scale basic chemicals and intermediates. In these reactors, the reaction takes place in the form a heterogeneous catalyst. In addition to the synthesis of valuable chemicals, fixed bed reactors have been increasingly used in recent years to treat harmful and toxic substances. The most common arrangement is the multi tubular fixed bed reactor, in which the catalyst is arranged in the tubes, and the heat carrier circulates externally around the tubes. Fixed bed reactor for industrial synthesis are generally operated in a stationary mode under constant operating conditions over prolonged production runs, and design therefore concentrates on achieving an optimum stationary operation. However, the non-stationary dynamic operation mode is also great importance for industrial operation control. 4.2.1 Reaction Taking Place on the Catalyst The processes taking place on the catalyst may formally be subdivided into the following separate steps: 1. Mass transfer of reactants from the main body of the fluid to the gross exterior surface of the catalyst particle. 2. Molecular diffusion /Knudsen flow of reactants from the exterior surface of the catalyst particle into the interior pore structure.
  37. 37. 36 3. Chemisorption of at least of the reactants on the catalyst surface. 4. Reaction of the surface 5. Desorption of absorbed species from the surface of the catalyst. 6. Transfer of products from the interior catalyst pores to the gross exterior surface of the catalyst by ordinary molecular diffusion/Knudsen flow. 7. Mass transfer of products from the exterior surface3 of the particle into the bulk of the fluid. 4.2.2 Catalyst Properties: Diameter of catalyst= 0.6mm Bulk density of catalyst=810kg/m3 Voidage=0.32 Surface area=45m2 /g 4.2.3 Steps for design: • Particles solid density • Void volume of the catalyst • Reaction rate • Weight of catalyst • Volume of catalyst bed • Pressure drop and height of bed • Diameter of shell and tubes • Number of tubes • Shell and tube side coefficients • Overall coefficient 4.2.4 Particle solid density: ρp=ρb/1-Ɛp ρb=810kg/m3 , Ɛp=0.32 ρp= 1191.2 kg/m3 4.2.5 Void Volume of catalyst: Vc=Ɛp/ρp = 0.268 g/cm3
  38. 38. 37 4.2.6 Reaction Rate: Kinetic study shows that in determining the rate of reaction only main reaction will participate while both side reactions can be neglected. If forward reaction constant is k1 and reverse reaction constant is k2 then Arrhenius Parameters are given: A1=6.5x105 A2=1.36x108 -rB=k1CB-k2CM 173.8moles/m3 hr 4.2.7 Weight of catalyst: W/F=∫dX/r Here W is the weight of catalyst needed for the required conversion which comes out: 2998 kg 4.2.8 Volume of the catalyst bed: Vb=W/ρb = 2998/810 3.70m3 4.2.9 Pressure drop in the bed: Pressure drop in packed bed is calculated by Ergun equation: ∆𝑃 𝐿 = 150𝜇(1 − 𝜀)2 𝑢 𝑜 𝜀3 𝑑 𝑝 2 + 1.75(1 − 𝜀)𝜌𝑢 𝑜 2 𝜀3 𝑑 𝑝 This gives the pressure drop per unit length of the bed as: 1060N/m3 4.2.10Height of the bed: The optimum value of pressure drop in the packed bed is almost 3 to 10% of its total pressure. So by iteration, we get Height of bed=4m 4.2.11Total cross section area of tubes: At=Vb/L =3.7/4
  39. 39. 38 =0.952m2 4.2.12Tube diameter: Tube diameter to particle ratio should not be very low Do = 0.11m , Thickness=0.005m So, Di =0.10m 4.2.13Tube area: A=∏Di 2 /4 =0.009m2 4.2.14Number of tubes: No. of tubes: Nt=0.952/0.009 =105 tubes 4.2.15Pitch of the tube: Pt=1.25Do =0.143m 4.2.16Bundle diameter: Db=Do(Nt)1/n /k1 n=2.14 and k1=0.319 =1.70m 4.2.17Shell Diameter: Shell diameter is taken approximately 5% more than bundle diameter, so; Ds=1.7+(.05x1.7) 1.79m 4.2.18Shell side heat transfer co efficient: Baffle spacing=0.3Ds=0.6m Area of cross flow=(Pt-Do)*Ds*Bs/Pt =0.218m2 Re=Gdeq/μ
  40. 40. 39 deq is measured by the expression: Deq = 4( pt 2 (0.86pt) − 1 2 πdo 2 4 πdo 2 Deq = 0.084m Pr=cpμ/k Re=GDeq/μ Re=2505 Pr=12.3 Assuming 35% baffle cut: Jr=0.13 Shell side coefficient is calculated by: hs=kjrRePr1/3 /de
  41. 41. 40 763.2W/m2 C 4.2.19Tube side co efficient: Cross sectional area= Πdi 2 /4 =0.009m2 Total flow area=0.009x105 =0.94m2 mass velocity=G=mass flow rate/flow area =2.05kg/sm2 Re=1000 Pr=12.3 Jh=0.02 hi=krhRePr0.33 /di
  42. 42. 41 378W/m2 C Correction for tube co efficient: ho=hiDi/Do 343W/m2 C 4.2.20Overall Heat transfer Co efficient: Overall coefficient= hohs/ho+hs 236.6W/m2 C 4.2.21Specification Sheet: Identification Catalyst weight 2998kg Volume 3.7m3 Height of bed 4m No. of tubes 105 Pitch 0.143m Bundle diameter 1.70m Tube side co efficient 343W/m2 C Shell side co efficient 763W/m2 C Overall co efficient 236.6 W/m2 C Baffle cuts 35% Baffle spacing 0.6m %age of pressure drop 4.2% Tube arrangement Triangular pitch Inlet temperature 326.3K Outlet temperature 374K
  43. 43. 42 4.3 DESIGN OF DISTILLATION COLUMN In industry it is common practice to separate a liquid mixture by distillating the components, which have lower boiling points when they are in pure condition from those having higher boiling points. This process is accomplished by partial vaporization and subsequent condensation. 4.3.1 CHOICE BETWEEN PLATE AND PACKED COLUMN Vapour liquid mass transfer operation may be carried either in plate column or packed column. These two types of operations are quite different. A selection scheme considering the factors under four headings. Factors that depend on the system i.e. scale, foaming, fouling factors, corrosive systems, heat evolution, pressure drop, liquid holdup. i) Factors that depend on the fluid flow moment. ii) Factors that depends upon the physical characteristics of the column and its internals i.e. maintenance, weight, side stream, size and cost. iii) Factors that depend upon mode of operation i.e. batch distillation, continuous distillation, turndown, intermittent distillation. The relative merits of plate over packed column are as follows: i) Plate column are designed to handle wide range of liquid flow rates without flooding. ii) If a system contains solid contents, it will be handled in plate column, because solid will accumulate in the voids, coating the packing materials and making it ineffective. iii) Dispersion difficulties are handled in plate column when flow rate of liquid are low as compared to gases. iv) For large column heights, weight of the packed column is more than plate column. v) If periodic cleaning is required, man holes will be provided for cleaning. In packed columns packing must be removed before cleaning. vi) For non-foaming systems the plate column is preferred. vii) Design information for plate column are more readily available and more reliable than that for packed column.
  44. 44. 43 viii) Inter stage cooling can be provide to remove heat of reaction or solution in plate column. ix) When temperature change is involved, packing may be damaged. For this particular process, “Benzene, Cumene, Propylene, Propane, DIPB ”, I have selected plate column because: i) System is non-foaming. ii) Temperature is high (90o C). 4.3.2 Choice of Plate Type There are four main tray types, the bubble cap, sieve tray, ballast or valve trays and the counter flow trays. I have selected sieve tray because: i) They are lighter in weight and less expensive. It is easier and cheaper to install. ii) Pressure drop is low as compared to bubble cap trays. iii) Peak efficiency is generally high. iv) Maintenance cost is reduced due to the ease of cleaning.
  45. 45. 44 4.3.3 DESIGNING STEPS OF DISTILLATION COLUMN  Calculation of Minimum Reflux Ratio Rm.  Calculation of Actual reflux ratio.  Calculation of theoretical number of stages.  Calculation of actual number of stages.  Calculation of diameter of the column.  Calculation of weeping point.  Calculation of pressure drop.  Calculation of thickness of the shell.  Calculation of the height of the column
  46. 46. 45 Component name Feed(mole fraction) Top(mole fraction) Bottom(mole fraction) i-butene 0.012 0.027 Methanol 0.0018 0.0386 0.00023 Mtbe 0.5706 0.092 0.938 Water 0.026 0.0211 0.03 DME 0.044 0.102 TBA 0.018 0.0318 n-butane 0.0037 0.00855 i-butane 0.3234 0.745 Total mass(kg/hr) 59625998 25884893 4.41*10^09 4.3.4 Design Data Item Distillation column Tray type Sieve tray No. of trays 12 Pressure 1 atm Height of column 7.306 m Diameter of column 0.729m Tray spacing 0.3048m Tray thickness 3mm Flooding 80% Hole area/Active area 0.10 m2
  47. 47. 46 Weir length 0.6806m Reflux ratio 0.168 Hole size 6 mm Down comer area 0.05004 m2 Hole area 0.030024 m2 Active area 0.30024 m2 Relative Volatilities () : Feed temperature = 337 K, Top temperature = 326K, bottom temperature = 380K Component Top relative volatility Bottom relative volatility Average relative volatility i-butane(LK) 7.699 5.084 6.3918 n-butane 5.65499 - 2.8274 i-butene 6.9487 - 3.4743 DME 14.1064 - 7.053 Mtbe(HK) 1 1 1 TBA - 0.1493 0.746123 Methanol 0.67395 0.99561 0.83478 Water 0.1524 0.2883 0.22037
  48. 48. 47 4.3.5 Rmin Using Underwood Method Using underwood equation q1 θα α θα α θα α C fCC B fBB A fAA       xxx As feed is at its bubble point so q = 1 By trial  = 1.7 Using equation of min. reflux ratio, 1R θα α θα α θα α m C dCC B dBB A dAA       xxx R min = 0.12 Rule Of Thumb: R = (1.2 ------- 1.5) Rm R = 1.4 Rmin R = 0.168 4.3.6 Calculating Minimum No. Of Stages Using Fenske’s equation      aveBC sX X X X m αlog dlog 1N B C C B          log(9.333) log N 0.00543 0.9557 0.0138 0.9564 m  = 10
  49. 49. 48 4.3.7 Theoretical Number Of Stages Using Erbar - Maddox graph (fig 11.11 C&R Vol. 6) N + partial condenser = 29 N = 28 4.3.8 Actual Number Of Stages: By rule of thumb tray efficiency=70% actual no.0f stages = theoretical no.of stages / tray efficiency = 42 stages 4.3.9 Determination Of The Column Diameter Top Conditions Bottom Conditions Ln = 9.868 Kgmol/hr Vn = 5 Kgmol/hr Average mol.Wt. = 76.506 Kg/Kgmol T = 100o C Liquid density = L = .769 gm/cm3 Vapour density = V = 4.31  10-3 gm/cm3 Lm = 86.74 Kgmol/hr Vm = 50.98 Kgmol/hr Average mol. wt. =120.1 Kg/Kgmol T = 179o C Liquid density = L = .715 g/cm3 Vapor density = V = 6.1  10-3 gm/cm3 Because liquid and vapour flow rates are greater at bottom so based upon bottom flow rates. 4.3.10 Vapour Load At Bottom 3600ρ V Q V m V   = 3600.0061 50.98  = 0 .278 m3 /Sec
  50. 50. 49 4.3.11Liquid Load at Bottom 3600ρ L Q L m L   = 3600.715 86.74  = 4  10-3 m3 /Sec 4.3.12Tray Dynamics i) Flow Parameter 0.5 L v m m LV ρ ρ V L F              = 0.5-3 .715 106.1 50.98 86.74              = 0.1572 FLV = Liquid Vapour Factor ii) Capacity Parameter Assumed tray spacing = 19 inches = 50 cm From Fig 3.23 (b) of “ Distillation Dynamics & Control”, sieve tray Flooding capacity, Csb(20) = 0.105 m/Sec Surface tension of system =  = 13.96 dynes/Cm Corrected Csb = Csb(20) 2.0 20        = .105 2.0 20 96.13       = .0877 m/Sec Now Uf = 0.5 V VL bs ρ ρρ C        = 0.5 .0061 .0061.715 .0877        = 0.945 m/Sec iii) Flooding Check Un = n V A Q = T V 0.88A Q = 417.88. 278.  = 0.7575 m/Sec
  51. 51. 50 Now F = 100 U Un       f =       945.0 .7575 100 = 79.15 iv) Calculation of Entrainment As FLV = 0.1572 and F = 79% From Figure 11.29 of coulson 6 , we calculate  = 0.005  = Fractional Entrainment factor since  < 0.1, so now process is satisfactory v) Column Diameter Let Flooding = 80% Un = Uf  F = 0.945  .8 = .756 m/Sec Un = flooding velocity based upon net area. Net area An = AT – Ad = 0.88 AT AT = 0.88 An = * n V 0.88U Q = 0.7560.88 .278  = 0.417 m2 AT = 4 π D2 D =  AT4 =  0.4174 D = 0.729 m = 2.39 ft.
  52. 52. 51 Now following information are available, Tower area = AT = 0.417 m2 Net area = An = 0.88 AT = 0.36696 m2 Active area = Aa = 0.76 AT = 0.30024 m2 Down comer area = Ad = 0.12 AT = 0.05004 m2 Hole area = Ah = 0.1 Aa = 0.030024 m2 vi) Tray Pressure Drop Ht =Hd + ( Hw + How ) + Hr a. Hw = 38.1 mm b. Dry Tray Drop Hd = 51 (Uh/Co)2 (V /L) Hd = Dry tray drop. Uh = Hole velocity = Qv/Ah Uh = h V A Q = 30024.00.1 0.278  = 9.254 m/sec Tray thickness = 3mm = 0.118 in. [For steel tray] Hole dia = 6 mm Aa Ah = 0.1 Using Fig. 11.34 of Coulson 6. We find “Co “ Co = Orifice Coefficient = 0.735
  53. 53. 52 Hd = 51 𝜌 𝑣 𝜌 𝐿 2 O h C U       = 51 0.0061 0.715 2 0.735 9.254       = 69.047 4.4 Liquid-liquid extraction unit design Liquid-liquid extraction has become an important separation technique in modern process technology. This is has resulted in the rapid development of a great variety of3extractor types, in the evaluation of which the chemical engineer must primarily depend on manufacturers’ literature. To the design, only three components that are considered, methanol, water and isobutene this is because of for most system containing more than four components, the display of equilibrium data and the computation of stages is very difficult. In such cases, the requirements are best obtained in the laboratory without detail study of the equalibria. Beside that for multicomponent separations also, special computer programs for these multistage operations embodying heat and material balances and sophisticated phase equilibrium relations are best left to professionals. Most such work is done by service organizations that specialize in chemical engineering process calculations or by specialize in chemical engineering organizations. Sieve tray (perforated plate) Column were choose for the extraction of these components. These multistage, countercurrent columns are very effective, both with respect to liquid-handling capacity and with respect to extraction efficiency, particularly for system of low interfacial tension, which do not require mechanical agitation for good dispersion. 4.4.1 Choice Of Solvent There is usually a wide choice of liquids to be used as solvent for extraction operations. It is unlikely that any particular liquid will exhibit all the properties considered desirable for extraction, and some compromise is usually necessary. The following factors need to be considered when selecting a suitable solvent for a given extraction – affinity for solute, partition ratio, density, miscibility, safety and cost. Based on the factors that need to be considered water was choosing as a solvent in this system.
  54. 54. 53 Physical properties from the literature Flowrate at the dispersed phase, QD = 1254.92 ft3/hr Flowrate at the continuous phase, Qc = 1.785 ft3/hr Density at the dispersed phase, ρD = 24.10 lb/ ft3 Density at the continuous phase ρC = 41.45 lb/ft3. The data was evaluated at system temperature and pressure. 4.4.2 Calculation of number of stages Extraction factor ε= 1 Number of theoretical stages = Nf = 𝑋𝑓−𝑌𝑠/𝑚 𝑋𝑟−𝑌𝑠/𝑚 -1 Where, Xf = 0.0195 Kg CH3OH/Kg Water Ys= 1.57×10-6 m= 0.002 By putting the values, Nf = 23 units In this equation, Xf= weight solute/weight feed solvent in the feed phase Ys = weight solute/weight extraction solvent in extract Xr = weight solute/weight feed solvent in the raffinate phase m= slope = Mass flow rate of solvent/ Mass flow rate of added solvent
  55. 55. 54 The number of mass transfer units is identical to the number of transfer units when extraction factor is one. So, Nor = 23 units By assuming column efficiency E to be 80% the number of real stages are, N= ( Nr-1)/E =(23-1)/0.8 N = 27 units 4.4.3 Sizing of the sieve tray Assume tray spacing h= 2ft h =4.5Vd 2 ρC/2gc∆ρ Where, Vd = Velocity of down comer ρC = density gc = gravitational constant So, Vd = 12471ft/hr Now, area of downcomer AD = QD / VD By putting the values AD comes out to be 0.1 ft2 Assume velocity through the holes Uh = 2880ft /hr. this large value is selected to avoid the formation of small droplets which decrease the column efficiency. So, total holes area = AHT = QD / UH The value of AHT comes out to be 0.43573ft2
  56. 56. 55 Tray area can be found by using the ratio of tray area AT to the holes area AH Tray area AT = 0.866 (ds)2 and holes area AH = 0.5(π / 4)(D0)2 ds and d0 are usually told by the manufacturer and their values usually range between do = 1/8 to 1/4 and ds varies from 0.75 to 0.85in. So, AT = 0.866 (0.75)2 = 0.48 in2 And AH = π / 8 (0.25)2 = 0.0245 in2 Tray area / hole area = 0.48 / 0.0245 So tray area AT = 19.88ft2 And the total tray area is = (tray area)(holes area) AT = 8.67 ft2 Number of holes AH = holes area = 0.5 ( π / 4)(do)2 Holes area = 0.0245 in2 Total number of holes NH = AHT/ AH = 0.43573(12)2 / 0.0245 = 2561 units Column parametes
  57. 57. 56 Number of trays = NT = Nr/ET = 23/0.8 = 29 trays Column height = ZT×NT = 2×29 CT = 58 ft Column and tray diameter are equal so, Column diameter = 3.32 ft Column area = 8.67 ft Net area = Ac-Ad = 8.67-0.1 = 8.57 ft2 Active area = Ac – 2Ad Aactive = 8.47 ft2 Now, height equivalent to theoretical stages (HETS) HETS = Column height/ Ideal stages = 58/23 = 2.52 Hor = column height/ Nor = 2.52 Weeping Evaluation
  58. 58. 57 Weir height = 50mm Hole diameter = 5mm Plate thickness = 5mm Weir crust can be found on the basis of maximum and minimum flow rates. Max how = {750 lwmax / ρD(lw)}2/3 = 29.73mm liquid Min how = {750 lwmin / ρD(lw)}2/3 = 7.17 mm liquid Where, Lw = weir length How = weir crest Low = liquid flow rate Pressure drop h = 51(uh / Co)2 (ρC / ρD) = 7.88 mm liquid Residual head hr hr = (12.5×103 ) / ρD = 32.89 mm liquid Total plate pressure drop = ht = 97.94 mm liquid Plate pressure drop ∆pt = 9.81×10-3 htρD
  59. 59. 58 = 365.147 Pa (N/m2 ) Provisional Plate Design Specification 1011MmColumn Diameter 27 TraysNumber of trays 0.6MTray spacing 5mMPlate thickness 9MTotal column height 7.88mm liquidPlate pressure drop SS304Plate material 9.3×10-3 m2 Downcomer area 0.805 m2 Column area 2.647 m2 Net area 0.773 m2 Active area 1.58×10-5 m2 Hole area 2560 m2 Number of holes 1.22705 unitsWeir length 50MWeir height
  60. 60. 59 4.5 Design of a distillation column: The design of a distillation column includes different steps which are as follows: • Calculation of Minimum number of stages.Nmin • Calculation of Minimum Reflux Ratio Rm. • Calculation of Actual Reflux Ratio. • Calculation of theoretical number of stages. • Calculation of actual number of stages. • Calculation of diameter of the column. • Calculation of pressure drop. • Calculation of the height of the column. 4.5.1 Calculation of Minimum no. of Plates The minimum no. of stages Nmin is obtained from Fenske equation which is, Nmin = ln[(xAD/xBD)(xBW /xAW)] ln (αAB) average
  61. 61. 60 Average geometric relative volatility = 2.277 So, Nmin = 10.76 ~ 11 4.5.2 Calculation of Minimum Reflux Ratio Using Underwood equations As feed is entering as saturated liquid so, q = 1 By trial,  = 1.75 Using equation of minimum reflux ratio, Putting all values we get, Rm = 1.407 4.5.3 Actual Reflux Ratio: The rule of thumb is: R = (1.1 ------- 1.5) R min So, taking R = 1.5 R min R = 2.11 4.5.4 Theoretical no. of Plates: Using Gilliland relation, From which the theoretical no. of stages to be, N= 20 1R θα α θα α m B DBB A DAA     xx q1 θα α θα α B fBB A fAA     xx                  566.0 minmin 1 175.0 1 R RR N NN
  62. 62. 61 4.5.5 Actual number of stages: By rule of thumb tray efficiency = 70% So, No. of actual trays = 20/0.70 = 29 4.5.6 Number of plates in rectification and stripping section: The Kirk bride method is used to determine the ratio of trays above and below the feed point. From which, Number of Plates in rectification section = ND = 8 Number of Plates in stripping section = NB = 21 • So, the feed plate is 9th from above 4.5.7 Diameter of column Using the following equation: FLV = Liquid Vapor Factor = 3.04 E -2 Assumed tray spacing = 18 inch (0.5 m) Csb = 0.29 Surface tension of Mixture = σ = 19 dynes/Cm Vnf = 2.23 m/sec Assume 90% of flooding then Vn=0.9Vnf So, actual vapor velocity, Vn = 2 m/sec Net column area used in separation is                              2 log206.log DHK BLK LK HK B D x x x x D B N N 0.5 L v n n LV ρ ρ V L F              5.02.0 20               v vl CV sbnf  
  63. 63. 62 An = mv/Vn Volumetric flow rate of vapors = mv mv = (mass vapor flow rate) /(3600)(vapor density) mv = 6.997 m3 /sec Now, net area An = mv/Vn = 3.13m2 Assume that down comer occupies 15% of cross sectional Area (Ac) of column thus: Ac = An + Ad Where, Ad = down comer area Ac = An + 0.15(Ac) Ac = An / 0.85 Ac= 3.68 m2 So Diameter of Column Is Ac =(π/4)D2 D = (4Ac/π)0.5 For the top area comes out to be D = 2.17 meter 4.5.8 Weir calculations Net Area (An) = Ac - Ad = 3.128 m2 Down comer area ‘Ad’ = 0.15Ac = 0.552 m2 From figure Lw / dc = 0.81 Lw = 2.17*0.81 = 1.76 m And by rule of thumb
  64. 64. 63 weir height ‘ℎ 𝑤′= 10 mm Hole diameter, dh = 5 mm Plate thickness = 5 mm 4.5.9 Pressure drop Vapor velocity through hole 𝑈0 = 𝑣𝑎𝑝𝑜𝑟 𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 ℎ𝑜𝑙𝑒 𝑎𝑟𝑒𝑎 Putting values 𝑈0 = 22.12 m/s Using equation Putting values, hd = 71.7 mm liquid Total pressure drop over plate is given by Where ℎ 𝑤 = 10mm ℎ 𝑜𝑤 = 43.4( 𝑞 𝑙 𝑊 )2/3 = 20.3mm So total ℎ 𝑡 = 102 mm ∆ P = 1.6 kPa per plate Area of 1 Hole = (π/4) Dhole 2 L V o h d C U h   2 ˆ 51          rowwdt hhhhh  )(
  65. 65. 64 = 0.00002 m2 Active area Aa=Ac-2Ad = 2.576 m2 Hole area Ah (take 10% Aa) = 0.1 × 2.576 Area of N Holes = 0.2576 m2 So, Number of Holes = 10288 4.5.10Height of column To calculate height of column, we use: Z= N. (T.S) Z= 29 * 20 = 580 in. Z = 14.72 m Where T.S = 20 inches (rule of thumb: tray spacing = 18in ---- 24 in) Here, Z = height of the column N= actual number of trays T.S= tray spacing Actual height of column Rule of thumb We add 1.2m for condenser at the top and 1.8m for reboiler at bottom So actual height of column = 17.7 m
  66. 66. 65 Selection of trays Three basic types of cross flow trays used are  Sieve Plate (Perforated Plate)  Bubble Cap Plates  Valve plates (floating cap plates) Why sieve plates? We prefer sieve plates because:  Pressure drop is low as compared to bubble cap trays  Their fundamentals are well established, entailing low risk.  The trays are low in cost relative to many other types of trays.  They can easily handle wide variations in flow rates.  They are lighter in weight.  Maintenance cost is reduced due to the ease of cleaning.
  67. 67. 66 4.5.11Specification sheet of distillation column: Identification Item Distillation column No. required 1 Column type Sieve Tray Mechanical design parameters Column diameter 2.2m Area of column 3.68m 2 Tray spacing 0.5m Hole diameter 5mm Weir height 10mm Tray thickness 5mm Thermal design parameters number of actual stages 29
  68. 68. 67 4.6 DESIGN OF HEAT EXCHANGER: 4.6.1 INTRODUCTION Shell and tube heat exchangers are the most versatile type of heat exchangers. They are used in process industries, in conventional and nuclear power stations as condensers, in steam generators in pressurized water reactor power plants, in feed water heaters and in some air conditioning and refrigeration systems. They are also proposed for many alternative energy applications including ocean, thermal and geothermal. Shell and tube heat exchangers provide relatively large ratios of heat transfer area to volume and weight and they can be easily cleaned. Shell and tube heat exchangers offer great flexibility to meet almost any service requirement. The reliable design methods and shop facilities are available for their successful design and construction. Shell and tube heat exchangers can be designed for high pressures relative to the environment and high pressure differences between the fluid streams. Shell and tube heat exchangers are built of round tubes mounted in a cylindrical shell with the tubes parallel to the shell. One fluid flows inside the tubes, while the other fluid flows across and along the axis of the exchanger. The major components of this exchanger are tubes (tube bundle), shell, front-end head, baffles and tube sheets. Shell types-various front and rear head types and shell types have been standardized by Tubular Exchanger manufacturers Association (TEMA). The E-shell is the most common due to its cheapness and simplicity. In this shell, the shell fluid enters at one end of the shell and leaves at the other end that is there is one pass on the shell side. The tubes may have a single or multiple passes and are supported by transverse baffles. This shell is the most common for single- phase shell fluid applications. With a single-tube pass, a nominal counter flow can be obtained. The design of a shell and tube heat exchanger is an iterative process because heat transfer coefficients and pressure drop depend on many geometric factors, including shell and tube diameters, tube length, tube layout, baffle type and spacing and the numbers of tube and shell passes, all of which are initially unknown and are determined as part of the design process. In production of MTBE, heat exchanger is very important equipment. Heat exchanger is used to increase or to decrease the mixture to the desired temperature. In order to make the process of production of MTBE taking place in the system, it is important to make the system at the correct environment. The heat exchanger that we used here is the shell and tube exchanger. Height of column 15m Reflux ratio 2.11 Pressure drop per tray 7.5 kPa
  69. 69. 68 Shell and tube heat exchanger is the most common type of heat exchanger used in the industry. This is because it has many advantages. The advantages are: -  It provided a large transfer area in a small space.  Good mechanical layout: a good shape for pressure operation.  Used well-established fabrication techniques  It can be constructed from a wide range of materials  It can be clean easily.  Well-established design procedures  Single phases, condensation or boiling can be accommodated in either the tubes or the shell, in vertical or horizontal positions.  Pressure range and pressure drop are virtually unlimited and can be adjusted independently for the two fluids.  Thermal stresses can be accommodated inexpensively.  A great variety of materials of construction can be used and may be different for the shell and tubes. The chemical engineering design for the heat exchanger is also known as thermal. The design requires the calculation of the heat transfer area required. From this value, design features of the unit such as the tube and shell size, tube counts and layout is determined. In addition, then pressure loss of the fluids across the unit is also calculated by determined the pumping capacity required. The calculation of the design is base on the first heat exchanger. The chemical design is based on Bell's and Kern method. Bell's method accounts for the major bypass and leakage streams. Kern method was based on experimental work on commercial exchangers with standard tolerances and will give a reasonably satisfactory prediction of heat-transfer coefficient for standard design. Properties of raw material (iso butane and n-butane) and steam for heat exchanger: Component Raw material (isobutane and n-butane) Steam Temperature inlet © t1=117 T1=350 Temperature outlet© t2=250 T2=350 Specific heat j/kgC 2155 2010 Thermal conductivity W/mK 0.07 0.03065
  70. 70. 69 Densitykg/m3 485 0.49375 Viscosity kg/ms 1.3E-04 1.55263E-05 Feed flow rate kg/s 10.9314 1.7649 Schematic Diagram
  71. 71. 70 4.6.2 Calculation of Heat Duty Heat duty= Enthalpy of leaving stream - Enthalpy of entering stream Heat duty = 8310 KW 4.6.3 Calculation of LMTD LMTD =157.23 C 𝑅 = T1−T2 t2−t1 , 𝑆 = t2−t1 T1−t1 R= 0 S = 0.57 4.6.4 Overall Heat Transfer Coefficient (U) Assume Overall Heat Transfer Coefficient (U) As: From table 12.1 (Coulson & Richardson’s Chemical Engineering), we take overall Coefficient U = 450 W/m2 C 4.6.5 Calculation of heat transfer Area 𝐴 = Q 𝑈∆𝑇 , A=117.45 m2 4.6.6 Number of Tubes Calculation From table 12.3 (Coulson & Richardson’s Chemical Engineering), we take standard pipe of: Inside diameter, di = 16 mm Outside diameter, do = 20 mm
  72. 72. 71 Length of pipe is assumed as 16 ft. Length, L = 4.88 m Area of the pipe can be calculated using equation below Area = LπD Area = (4.88)(3.142)(0.02) Area = 0.3067 m2 4.6.7 Number of Tubes Nt=A/a Nt =383 4.6.8 Bundle and Shell Diameter Calculation The triangular pitch of 1.25 is chosen as the tube arrangement. Db=Do(Nt)1/n /k1 From table 12.4 (Coulson & Richardson’s Chemical Engineering), for 1.25 triangular pitch, number of passes = 2, then we can obtain K1 = 0.249 n1 = 2.207 Db=0.556m Assume using pull-through floating head type. From figure 12.10 (Coulson & Richardson’s Chemical Engineering), for bundle diameter 0.33, bundle clearance is 93 mm. Shell diameter,
  73. 73. 72 Ds=0.556+0.093=0.649m 4.6.9 Tube Side Coefficient , hi T mean = 183.5 C Tube / Pass = 383/2 = 191.5 Total flow rate area = (cross sectional area)(tube / pass ) Total flow rate area = .0383 sq.meter. Steam mass velocity, Gt= (steam flow rate)/(total flow rate area) Gt= 46.08 Kg /sm2 Steam linear velocity, u1= (Gt)/(steam density) U1 = 93.328 kg /ms L/di=4.88/0.016=305 4.6.10Reynolds And Prandtel Number Re=47486.6 Jh=4.1E-03 (from Coulson and Richardson) Pr=Cpμ/k =1.0180 Tube side co efficient can be calculated by: hi=kjhRePr0.33 /di =320.34 W/mC
  74. 74. 73 4.6.11Shell Side Coefficient, hs lB= Ds/ 3 = .2163 Tube pitch, pt= 1.25 do = (1.25)(0.02) = 0.025 m Flow area, (As) = (Pt−do)(Ds)(lB) Pt As = 0.028 m2 Mass velocity=Gs= 𝑊𝑠 𝐴𝑠 Gs = 390.4 kg / m2 Shell side velocity Us = Gs ρ Us = 0.8049 m/s 4.6.12Shell Side Equivalent Diameter for Triangular Pitch Arrangement 𝐷𝑒 = 1.10 𝑑𝑜 (pt 2 - 0.917do2 ) = 0.0142m Reynolds number Re=Gsde/μ Re = 42642.8 Prandtl Number Pr=cpμ/k
  75. 75. 74 Pr = 4.0023 4.6.13Shell Side Coefficient, hs Choose baffle cut of 25%, from figure 12.30 (Coulson & Richardson’s Chemical Engineering),can obtained: Jf = 2.7 * 10 -2 Assumed that the viscosity correction is negligible hs=kjfRePr1/3 /de hs = 20015.9 W / m2 0 c 4.6.14Overall Heat Transfer Coefficient, Uo Material of construction = carbon Steel Thermal conductivity of the tube wall Kw = 38 W/mC Assumed dirt coefficient as hid = 8500 W/m2 C hod = 8500 W/m2 C U = 426.33 W / m2 o C 1 Uo = 1 hs + 1 hod + doln( do di ) 2kw + 1( do di ) hid + 1( do di ) hi
  76. 76. 75 4.6.15Tube Side Pressure Drop Reynolds number = Re = 47486.6 Jf = 3.1 * 10 – 3 Neglect the viscosity correction term ∆Pt=Np[8jf(L/di)(μ/μw)-m +2.5] ρiμi 2 /2 ∆Pt=43.281 kPa 4.6.16Shell Side Pressure Drop Reynolds number = 42642.8 Jf = 2.7 * 10 – 2 Shell side pressure drop can be calculated using equation below ` ∆P=8jf(Dd/de)(L/lB)(ρμs/2)(μ/μw)-0.14 ∆P=23.473 kPa 4.6.17Specification Sheet Unit : Exchanger-100 Fluid allocation SHELL SIDE TUBE SIDE Fluid name STREAM 1 STREAM 2 Fluid quantity (kg/hr) 6353.64 39353 Temperature (In/Out) o C 350 350 117 250 Density Vapor/Liquid kg/m³ 0.49375 485 Viscosity, kg/ms 1.55263 x 10-5 1.30005 x 10-4 Feed flowrate, kg/s 1.7649 10.9314
  77. 77. 76 Specific heat, j/kgoC 2010 2155 Thermal conductivity, W/mK 0.0306575 0.07 Molecular wt, Vap 28,23 Pressure Drop (Kpa) 23.247 43.281 Heat Transfer Coefficient (W/m2 C) 20015.9 320.34 Fouling Resist. (min) m² K/W 0 0 Velocity m/s 0.8049 5 HYSYS 5.1 The Design Based On Hysys Simulation There are two method that was used in calculating the mass balance and energy balance for the process which is: i) Manual calculation ii) Hysys simulation Hysys program was used to see whether the design could be run or not. Using Hysys the calculation of the process was calculated automatically when the parameter that needed was insert. Then if the parameter that was insert is logic so Hysys program can calculated the result and the equipment can converge. If the data that was inserted was illogical the equipment cannot converge and the calculation cannot be done. Here the simulation of heat exchanger has done on ASPEN HYSYS. For this purpose following steps are followed:  Enlist the components
  78. 78. 77  Select the fluid package. Here we used Peng Robinson because all the components are hydrocarbons
  79. 79. 78  Then enter into simulation environment and select the shell and tube heat exchanger from the object pellet  Give the operating conditions to the stream for both the shell side and tube side  Specify the steams and give the pressure drop to the exchanger
  80. 80. 79  Run the simulation  Results are comparable 5.2 Results Manual Calculations HYSYS Calculations Heat Exchanger E 100 E100 Surface are, A (m2 ) 117.45 125.5 Overall heat transfer 426.33 145.98 Tube side pressure drop, (kPa) 43.281 45.56
  81. 81. 80 Shell side pressure drop ,(k Pa) 23.247 45.26 Number of tube, Nt 383 383 Duty, Q (Kw) 8310 8390 Results are approximately similar. 6 ENVIRONMENTAL IMPACT ASSESSMENT 6.1 INTRODUCTION An important procedure for ensuring that the likely effects of new development on the environment are fully understood and taken into account before the development is allowed to go ahead” EIA is required to be a full-disclosure statement. This includes project parameters that will have a positive environmental effect, negative impact, or no impact whatsoever. Generally, design engineers will only be involved with a small portion of the EIA preparation, in accordance with their expertise. However, each individual should be aware of the total scope of work necessary to prepare the EIA as well as the division of work. This will minimize costly duplication, as well as provide the opportunity for developing feasible design alternatives. The preparation of an Environmental Impact Assessment requires determining what environmental standards require compliance by the project, obtaining baseline data, examining existing data to determine environmental safety of the project, preparing an effluent and emission summary with possible alternatives to meet acceptable standards, and finally preparing the environmental statement or report. Since it may require a full year to obtain baseline data such as air quality, water quality,
  82. 82. 81 ambient noise levels, ecological studies, and social surveys, emissions and effluents, studies should take place concurrently to avoid delay in preparing the EIA. The emissions and effluents studies must include all “significant” sources of pollution. Omission of data could cause inconsistencies that could result in further time delays when negotiating with the regulatory agencies issuing the many required construction permits. It becomes clear that environmental considerations not only can play a major factor in the choice of selecting a plant site but can also be quite costly. The American Petroleum Institute has estimated that the preparation of an EIA for each site considered may range from $50,000 for small projects to $1.5 million for a large petroleum refinery. On the other hand, a detailed environmental assessment may quickly eliminate possible sites because of their highly restrictive standards. Nowadays, environmental issues become very important. Besides this, a good waste treatment system is also important in order to reduce and minimize environmental pollutants. The chemical waste in the form of solid, liquid and gases must be treated before being discharged into sewage, drain and atmospheres. Any chemical plant to be set up in Malaysia must follow the rules and regulations under the Department of Environment (DOE) Malaysia, which includes the Environmental Quality Act 1974. Under Environmental Quality Act (Sewage and Industrial Effluents) Regulation 1979 and Environment Quality Act (Clean Air) 1978. The plant owner or waste generator must ensure that waste generated disposed appropriately to prevent environmental pollution. The proper and suitable method should be implemented in dealing with the waste disposal. As our plant produces MTBE and other byproducts like raffinate but generally they are not hazardous to the environment and human if safety measures are taken into consideration. These environmental considerations depend on the location of our plant. The plant will follow the Standard B of water quality measurement and also need some waste treatment facilities to minimize the pollution from our plant.
  83. 83. 82 6.2 Steps of EIA: 6.3 Impacts of project on environment: 6.3.1 STACK GASES 6.3.1.1 Gas Emission Treatment Direct flame combustion was used to burn the excess gas. Flare is usually open ended combustion unit. Therefore, the combustion process will be controlled by flow rate of gases mixture to prevent incomplete combustion. Another treatment is thermal combustion. It is an incinerator used in the cases where the concentration of combustible gases is too low to make direct flame incineration insufficient condition. The temperature of operation depends upon the type of pollutant in waste gas. Thermal combustion must be carefully designed to provide safe, efficient operation and to prevent incomplete combustion. Time, temperature, and oxygen must be carefully monitored. (Howard et. al 1985)
  84. 84. 83 Stack gas means the product of combustion process usually occur at machine or generator. It is usually the fuels used occurred in the complete combustion process, but it produced unwanted gas such as carbon monoxide, sulphur oxide and other gases. In our MTBE plant, the stack gases is only Hydrogen and it is stored in a special tank before being sold to interested company at market price. 6.3.2 WASTEWATER TREATMENT 6.3.2.1 Wastewater Characteristics Wastewater characteristics vary widely from industry to industry. Obviously, the specific characteristics will affect the treatment techniques chosen for use in meeting discharge requirements. Because of the large number of pollutant substances, wastewater characteristics are not usually considered on a substance-by-substance basis. Rather, substances of similar pollution effects are grouped together into classes of pollutants or characteristics are indicated below. Priority Pollutants Recently, greatest concern has been for this class of substances for the reasons given previously. These materials are treated on an individual-substance basis for regulatory control. Thus each industry could receive a discharge permit that lists an acceptable level for each priority pollutant. Organics The organic composition of industrial wastes varies widely, primarily due to the different raw materials used by each specific industry. These organics include proteins, carbohydrates, fats and oils, petrochemicals, solvents, pharmaceutical, small and large molecules, solids, and liquids. Another compilation is that a typical industry produces many diverse waste streams. Good practice is to conduct a material balance throughout an entire production facility. This survey should include a flow diagram, location and sizes of piping, tanks, and flow volumes, as well as an analysis of each stream. An important measure of the waste organic strength is the 5-day biochemical oxygen demand (BOD). As this test measures the demand for oxygen in the water environment caused by
  85. 85. 84 organics released by industry and municipalities, it has been the primary parameter in determining the strength and effects of a pollutant. This test determines the oxygen demand of a waste exposed to biological organisms (controlled seed) for an incubation period of five days. Usually this demand is caused by degradation of organics according to the following simplified equation, but reduced inorganics in some industries may also cause demand (i.e., Fe2+, S2- and SO32-). Organic waste + O2 CO2 +H2O In general, low-molecular-weight water-soluble organics are biodegraded readily. As organic complexity increases, solubility and bio-degradability decrease. Soluble organics are metabolized more easily than insoluble organics. Complex carbohydrates, proteins and fats and oils must be hydrolyzed to simple sugars, amines, and other organics acids prior to metabolism. Petrochemicals, pulp and paper, slaughterhouse, brewery, and numerous other industrial wastes containing complex organics have been satisfactorily treated biologically, but proper testing and evaluation is necessary. Inorganics The inorganics is most industrial wastes are the direct result or inorganic compounds in the carriage water. Soft-water sources will have lower inorganics than hard-water or saltwater sources. However, some industrial wastewaters can contain significant quantities of inorganics which result from chemical additions during plant operation. Many food processing wastewaters are high in sodium. While domestic wastewaters have a balance in organics and inorganics, many process wastewaters from industry are deficient in specific inorganic compounds. Biodegradation of organic compounds requires adequate nitrogen, phosphorus, iron, and trace salts. Ammonium salts or nitrate salts can provide the nitrogen, while phosphates supply the phosphorus. pH and Alkalinity
  86. 86. 85 Wastewaters should have pH values between 6 and 9 for minimum impact on the environment. Wastewaters with pH values less than 6 will tend to be corrosive as a result of the excess hydrogen ions. On the other hand, raising the pH above 9 will cause some of the metal ions to precipitate as carbonates or as hydroxides at higher pH levels. Alkalinity is important in keeping Ph values at the right levels. Bicarbonate alkalinity is the primary buffer in wastewaters. It is important to have adequate alkalinity to neutralize the acid waste components as well as those formed by partial metabolism or organics. Many neutral organics such as carbohydrates, aldehydes, ketones, and alcohols are biodegraded through organics acids which must be neutralized by the available alkalinity. If alkalinity is inadequate, sodium carbonate is a better form to add than lime. Lime tends to be hard to control accurately and results in high pH levels and precipitation of the calcium which forms part of the alkalinity. In a few instances, sodium bicarbonate may be the best source of alkalinity. Temperature Most industrial wastes tend to be on the warm side. For the most part, temperature is not a critical issue below 37⁰C if wastewaters are to receive biological treatment. It is possible to operate biological wastewater-treatment systems up to 65⁰C with acclimated microbes. Low- temperature operations in northern climates can result in very low winter temperatures and slow reaction rates for both biological treatment systems and chemical treatment systems. Increased viscosity of wastewaters at low temperatures makes solid separation more difficult. Increased viscosity of wastewaters at low temperatures makes solid separation more difficult. Efforts are generally made to keep operating temperatures between 10 and 30⁰C if possible. 6.3.2.2 Liquid Waste Treatment Equalization Treatment Liquid treatment generally is necessary in any plant. In our plant, we also have liquid treatment but in general, we only state the general method, as our plant does not produce any significant liquid waste. In any liquid waste treatment, we need equalization treatment. The equalization treatment is an initial procedure in liquid waste treatment. The purpose of equalization is to minimize and control the fluctuation in liquid waste characteristic. Besides it provides the
  87. 87. 86 suitable and optimum condition for biological and chemical treatment. It also provides adequate damping to minimize the chemical consumption. The procedure will occur in the equalization tank. The size of tank and time of equalization process depend on the liquid waste amount. The Activated Sludge process will be used for this treatment. It is carried out in Aerobic condition. The main purpose of activated sludge process is to remove soluble and insoluble organic matter that converted into flocculants microbial suspension and settable microbial. It also permits the use of gravitational solid liquid separation technique for the above requirement. The organic matter where measured in the form of BOD and COD serves as food and energy source for microbial growth. It converts the pollutant into microbial cell and oxidized end product such as CO2 and H2O by microbial activities. Therefore, Submersible Aerator as mixing device will supply the oxygen and nutrient into aeration tank and therefore improves the quality of the liquid. (Howard et. al, 1985) 6.3.2.3 Solid Waste Treatment The solid waste treatment will be minimized by regenerating the catalyst. Regeneration processes depend on the characteristic of catalyst after whole reaction. Licensed contractor will dispose the solid waste to follow the DOE regulation. By the way, the scheduled maintenance activities will be implemented. Dewatering system will be used to solidify and extract the catalyst. Therefore, clarifier and filter press were used in these treatments. Clarifier is used to clarify any impurities before going through the filters. The size of equipment depends on the flow rate and holding time of these processes. Maintenance activities will be scheduled based on the availability of workers and machines. Skilled and experienced workers will do the maintenance activities, (Bailed, 1995). 6.3.3 Waste Minimization Waste minimization means the optimization process to minimize the waste come out of the plant. It will be done by source reduction and recovery of the sources. The source reduction refers to preventative measured taken to reduce the amount of waste, which produced in this process. Recovery of the sources is aimed to reuse the excess methanol to produce the MTBE. Waste production from the plant could be reduced by implementing these procedures:
  88. 88. 87  Raw material modification,  Product reformulation,  Process modification,  Improvement in operating practices. The most important is by improving the product yield and this means minimization of waste generation. It will be accomplished through improvement in catalyst efficiency and proper maintenance activities. 7 HAZOP STUDY 7.1 Introduction A HAZOP survey is one of the most common and widely accepted methods of systematic qualitative hazard analysis. It is used for both new or existing facilities and can be applied to a whole plant, a production unit, or a piece of equipment It uses as its database the usual sort of plant and process information and relies on the judgment of engineering and safety experts in the areas with which they are most familiar. The end result is, therefore reliable in terms of engineering and operational expectations, but it is not quantitative and may not consider the consequences of complex sequences of human errors. The objectives of a HAZOP study can be summarized as follows: To identify the areas of the design that may possess a significant hazard potential.  To identify and study features of the design that influence the probability of a hazardous incident occurring.  To familiarize the study team with the design information available.  To ensure that a systematic study is made of the areas of significant hazard potential.  To identify pertinent design information not currently available to the team.  To provide a mechanism for feedback to the client of the study team's detailed comments. 7.2 Steps Conducted In Hazop Study  Specify the purpose, objective, and scope of the study. The purpose may be the analysis of a yet to be built plant or a review of the risk of unexisting unit. Given the purpose and the circumstances of the study, the objectives listed above can he made more specific. The scope of the study is the boundaries of the physical unit, and also the range of events and variables
  89. 89. 88 considered. For example, at one time HAZOP's were mainly focused on fire and explosion endpoints, while now the scope usually includes toxic release, offensive odor, and environmental end-points. The initial establishment of purpose, objectives, and scope is very important and should be precisely set down so that it will be clear, now and in the future, what was and was not included in the study. These decisions need to be made by an appropriate level of responsible management.  Select the HAZOP study team. The team leader should be skilled in HAZOP and in interpersonal techniques to facilitate successful group interaction. As many other experts should be included in the team to cover all aspects of design, operation, process chemistry, and safety. The team leader should instruct the team in the HAZOP procedure and should emphasize that the end objective of a HAZOP survey is hazard identification; solutions to problems are a separate effort.  Collect data. Theodore16 has listed the following materials that are usually needed.  Process description.  Process flow sheets.  Data on the chemical, physical and toxicological properties of all raw materials,, intermediates, and products. Piping and instrument diagrams (P&IDs).  Equipment, piping, and instrument specifications.  Process control logic diagrams.  Layout drawings.  Operating procedures.  Maintenance procedures.  Emergency response procedures.  Safety and training manuals. 7.3 HAZOP METHOD FLOW DIAGRAM A HAZOP study is conducted in the following steps:
  90. 90. 89 7.4 HAZOP Guide Words and Meanings Guide Words Meaning

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