Chalmers University of Technology
Production of MEK Preliminary Plant Design Aina Viladot, Pilar Pizones, Ananda Subramani Kannan Group 9
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Table of Contents 1. INTRODUCTION .............................................................................................................................. 4 2. Reactor Design .................................................................................................................................... 5 3. SEPARATION PROCESS: ................................................................................................................ 7 3.1. CONDENSER: ................................................................................................................................ 7 3.2. ABSORPTION COLUMN ................................................................................................................ 7 3.3. EXTRACTION COLUMN ................................................................................................................. 8 3.4. DISTILLATION COLUMN 1 ............................................................................................................. 8 3.5. DISTILLATION COLUMN 2 ............................................................................................................. 9 4. EQUIPMENT SIZING ..................................................................................................................... 11 4.1 Extraction column sizing ............................................................................................................. 11 4.2 Absorption and Distillation sizing ............................................................................................... 12 Flash vessel sizing .............................................................................................................................. 12 5. Energy Analysis ................................................................................................................................ 13 5.1 Base Case .................................................................................................................................... 13 5. 2 Process Integration Assessments ............................................................................................... 16 5.2.1 Separation MER - ................................................................................................................. 16 5.2.2 Capital - Energy Trade-offs ................................................................................................... 21 5.3 Heat exchanger Network design ................................................................................................. 22 5.3.1 Practical ramifications of the heat exchanger network ....................................................... 23 5.4 Design of reactor heating ............................................................................................................ 24 5.4.1 Series Heating ...................................................................................................................... 24 5.4.2 Parallel Heating – ................................................................................................................. 26 6. Process Economics............................................................................................................................ 27 6.1 Cost estimations.......................................................................................................................... 27 6.1.1 Capital Cost Estimates .......................................................................................................... 27 6.1.2 Operating cost estimates ......................................................................................................... 32 7. HIGH PRESSURE STEAM GENERATION................................................................................... 38 7.1 Methodology ............................................................................................................................... 38 7.1.1 Base Case – High Pressure Steam Generation: Parallel Heating ......................................... 39 7.1.2 Practical Ramifications of the background-foreground analysis ......................................... 40 7.2 Economic Evaluation – ................................................................................................................ 40 7.2.1 Electricity vs. Steam ............................................................................................................. 40 7.2.2 Augmenting electricity production ...................................................................................... 41 1
9 Conclusions ........................................................................................................................................ 43 10 References ....................................................................................................................................... 44 11. APPENDIX ..................................................................................................................................... 45 11.1 APPENDIX 1: MASS BALANCE .................................................................................................... 45 11.2 APPENDIX 2. Economics ............................................................................................................ 48 11.2.1 Equipment: ......................................................................................................................... 48 11.2.1.1 Absorption tower ............................................................................................................ 48 11.2.1.2 Distillation column 1 ....................................................................................................... 48 11.2.1.3 Distillation column 2 ....................................................................................................... 49 11.2.1.4 Extraction column ........................................................................................................... 49 11.2.1.5 Flash
11.2.1.6 Pumps ............................... 50
11.2.2 Operating cost ........................................................................................................................ 50 11.2.3 Summary costs and revenues (NPW and NFW): .................................................................... 51 11.3 APPENDIX 3: Sizing .................................................................................................................... 54 11.3.1. Extraction column sizing ................................................................................................... 54 Calculation of cross-sectional area ............................................................................................... 54 11.3.2Calculation of the extraction column height. ..................................................................... 56 11.3.3. Flash vessel sizing .............................................................................................................. 57 11.4 APPENDIX 4: Heat Exchanger Network Design ......................................................................... 59 11.5 APPENDIX 5: Flow sheet ............................................................................................................ 61 11.6 APPENDIX 6: heat exchanger area estimation .............................................................................. 0
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Abstract This project will be focused on the preliminary design of the process plant for methyl ethyl ketone (MEK) production. It takes place at the chemical company Utopia. In order to investigate the possibility of manufacturing this product at its own site, as MEK is normally purchased. The proposed plant design can be considered to be a very profitable investment. The design would indicate that, the required production of 90000 tons/year at a reactor conversion of 93% is met with ease. The energy analysis on the other hand, is also indicative of very budding trends. There have been significant energy savings obtained with the process integration methods proposed. The heating media for the reactor were also chosen purely on the basis of the energy analysis. The parallel heating module would be recommended due to the lower consumption of natural gas in this configuration. Keywords – MEK,Process Integration, Heating media
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1. INTRODUCTION This project will be focused on the preliminary design of the process plant for methyl ethyl ketone (MEK) production. It takes place at the chemical company Utopia. In order to investigate the possibility of manufacturing this product at its own site, as MEK is normally purchased. It includes the reactor, separation process, process integration (Energy), and economy (investment cost estimate), as well as energy analysis, process design improvements, heat exchanger network, and design of utility system. The raw material for MEK production is 2-butanol, which is an intermediate product in one of the production lines. The 2-butanol is pumped from an existing tank to a preheater, and is then evaporated by heatexchanging with the reaction products. The vapor is then superheated to the reaction temperature, by using combustion gases from the reactor. The combustion gases are produced in a gas fired combustor. The superheated 2-butanol is fed to the reactor at a temperature between 400 and 500°C, and is reacting forming MEK as one of the reaction products. The MEK obtained is cooled in different stages and is then partially condensed. The liquid phase proceeds to a separation stage for recycling of 2-butanol and purification of MEK, while the gas phase is passed through an absorption column. The absorbed solution is lead to an extraction column for regeneration of the solvent, which is the returned to the absorption column, while the extract is forwarded to a distillation column for reprocessing. The 2-Butanol/MEK stream from this distillation column is conducted to the separation stage for MEK purification mentioned above.
Figure 1.1 Flowsheet for the production of MEK
Flowsheet for the production of MEK
Water H2 (MEK)
TCE
Water MEK
2-Butanol feed
Pump
Pre- Heater
Evaporator
Cooler
Superheater
Cooler
Absorption Column
Condenser
MEK
Extraction column
Plug Flow Reactor
MEK 99,5 wt%
TCE MEK
Distillation 1 MEK
2-But recycled
Distillation 2
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2. Reactor Design After the superheating, the 2-butanol is feed to the reactor at a temperature of 450°C, and it reacts forming MEK as one of the reaction products:
This is an endothermic and reversible reaction, with the following reaction ratio:
where pA, pk and pH are partial pressures in bar and the kinetics constants are:
The reactor chosen will be a multi-tube reactor which have parallel tubes with the same diameter and length, and has exact duplication of reaction conditions at scaling up. In the chemical industry, this type of reactor is used for catalytic gas-phase reactions over fixed-bed catalysts. It is very useful in exothermic and endothermic reactions because of the heating or cooling on the outside of the tubes (shell). In the proposed case, the gases from the combustion of the hydrogen will supply enough heat to the endothermic reaction and to the superheating of the 2-butanol. This is represented in the figure below.
Figure 2.1 Reactor
This is a more efficient way to achieve the temperature profile needed for the specifications of the final product. 5
To calculate the dimensions of the reactor, first of all we have to look at the operation conditions: The flux (G) in the reactor is based on the condition that:
where dp is the diameter of the particle (0,0032 m), µ is the viscosity of the gas mixture (1,89 · 10 -5 Pa.s), and ƐB the porosity of the bed (0,39). With this condition, interfacial gradients are avoided and a low pressure drop is maintained. This pressure drop is due to the porosity (ƐB) of the bed. The form of the catalyst particle would be a solid cylinder, so the pressure drop will be ∆P/L = 0,5 bar/m. The ratio dt/dp (diameter of the particle/ diameter of the tube) is given as:
This would yield dt = 0,037 m, (in order not to have wall effects.) The number of tubes required is given as: ̇
nt = 894 tubes. Finally we have to iterate with the length of the reactor in order to obtain a 93 % of conversion, which is our specification. L = 1.413 m The temperature profile within the reactor would have the behavior as shown in the figure. It is very essential that the heating utility profile is noted in the figure below. It is this utility profile that is used to determine the reactor heating options.
Figure 2.2 Temperatures profile
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3. SEPARATION PROCESS: 3.1. CONDENSER: We have the following data of the stream after the reactor: Table 3.1 Reactor product
REACTOR PRODUCT 405°C TEMPERATURE 179,3 KPa PRESSURE 324,3 Kmol/h MOLAR FLOW 2-but 3,62 MEK 48,21 COMPOSITION (molar %) H2 48,17 This stream is cooled, and its pressure is decreased in order to have it partially condensed to enter in the condenser. Then, before entering in the condenser, this stream will have a temperature of 33,61 °C and a pressure of 100 KPa, and it will be partially condensed. In the condenser, MEK and 2-butanol will be separated from the H2. The hydrogen will go in the top with traces of MEK, but the most part of the MEK will go in the bottom with the 2-butanol, and will be taken to the distillation column 2, the last step of the process.
3.2. ABSORPTION COLUMN The gas stream from the condenser has the following conditions: Table 3.2 Reactor product
REACTOR PRODUCT 33,61°C TEMPERATURE 100 KPa PRESSURE 187,9 Kmol/h MOLAR FLOW 2-but 0,39 MEK 16,47 COMPOSITION (molar %) H2 83,14 The stream entering the absorption column is composed of MEK and H2, (although there is a little 2butanol that wasn’t separated in the condenser). MEK will be absorbed by pure water at 25°C. In the column, there would be 13 stages with a constant pressure of 100 KPa. A minimum of the 95 % of the MEK in the feed has to be absorbed by the water, in order to satisfy the given restrictions. The hydrogen recovered will be sent to a burner to obtain energy. This gas phase out from the absorption column will have the following conditions that are very similar to the mass balances (stream 4): Table 3.3 Gas Phase from absorption column
GAS PHASE FROM ABS.COLUMN 25,51°C TEMPERATURE 100 KPa PRESSURE 163,6 Kmol/h MOLAR FLOW MEK 0,89 1,46 kmol/h COMPOSITION (molar %) H2 95,46 156,16 kmol/h 7
Water
3,26
5,33 kmol/h
3.3. EXTRACTION COLUMN The stream entering the extraction column has the following composition: Table 3.4 Stream to the extraction column
STREAM TO THE EXTRACTION COLUMN 34,74°C TEMPERATURE 100 KPa PRESSURE 1062 Kmol/h MOLAR FLOW 2-but 0,07 MEK 2,79 COMPOSITION (molar %) Water 97,14
Now, in the extraction column MEK is separated from water. Here, the separating agent used is 1,1,2Trocloroethane (TCE). This solvent is not very harmful and it’s a good choice to separate water from MEK. Moreover, TCE is much heavier than MEK (so it would be separated in a distillation column later). The TCE is feed in the top of the column (although in our HYSYS simulation is feed in the bottom). After the extraction, the water will be recirculated to the absorption column. To perform this, 10 stages and a pressure of 100 KPa will be needed. There will be no water in the bottom stream, it must be completely separated from the MEK, but the fraction of MEK in this stream won’t be more than a 40 mass%. This is because the solubility is not complete, as we can see in a phase equilibrium diagram for MEK/water/TCE.
3.4. DISTILLATION COLUMN 1 Before the distillation column 1, we will have a feed preheater, in order to get the feeding as saturated liquid. We will increase the pressure as well, so we can have with these two methods a better efficiency. We can see this when we do the energy analysis: If we shift the pressure in distillation column one, we will have a better process integration with major energy savings. With all this, we can see in the following table the conditions for the stream entering in the distillation column. Table 3.5 After feed preheating 1
AFTER FEED PREHEATING 1 93,46°C TEMPERATURE 265 KPa PRESSURE 53,67 Kmol/h MOLAR FLOW 2-but 1,37 MEK 54,96 COMPOSITION (molar %) Water 43,67
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In the first distillation column we want to separate TCE from MEK, with a purity as high as possible, with not more than 1 % of MEK on the bottom and not more than 1% of TCE on the top. The purity in the bottom is to prevent MEK from disturbing the extraction convergence. This stream of TCE is almost pure, and will be recycled after being cooled. In addition, we will not let TCE spread too much into the 2-butanol recycle to the reactor. The performance of the distillation column will depends on many factors. We can observe that if we increase the number of stages, or the reflux ratio, the separation will be better, but this will be more expensive: A large column will require more material, and increasing the reflux ratio means higher reboiler duty, so we have to look for a tradeoff among these factors. This first distillation column will work with 35 stages and a pressure of 265 KPa, with the stage 13 as inlet in order to achieve a better temperature profile, which can be seen in the figure.
Figure 3.1 temperature profile
The reflux ratio will be 1,25 times the minimum reflux ratio. A smooth profile, as above, would indicate that the positioning of the inlet feed is correct. After this distillation column we must have a valve in order to decrease the pressure and mixed the MEK stream with the stream coming from the condenser (MEK and 2-butanol)
3.5. DISTILLATION COLUMN 2 In this case we will have a feed preheater too, in order to get the feeding in its saturation point. The feeding, like we said before, is composed by the stream coming from the first distillation column and the stream coming from the condenser. The conditions and composition of the resulting stream is showed in the following table: Table 3.6 After heat preheating 2
AFTER FEED PREHEATING 2 80,33°C TEMPERATURE 100 KPa PRESSURE 166,6 Kmol/h MOLAR FLOW 2-but 7,04 COMPOSITION (molar %) MEK 92,96 In this second column we will obtain the final product, which must have 99 mass% of MEK (purity). The 2-butanol obtained in the bottom is recycled to the beginning of the process. We have to notice that small amounts of TCE that wasn’t removed in the distillation one would be present. This amount of TCE will be recycled with the 2-butanol to the beginning of the process.
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In this case, the distillation column will work with 23 stages, and with a pressure of 100 KPa. The inlet stage will be the 12, and we will have a reflux ratio of 1,58. As mentioned earlier we have had to find an optimum point between the number of stages and the reflux ratio, and among other factors like the conditions of the feeding.
The final product compositions are:
Table 3.7 Final product
FINAL PRODUCT TEMPERATURE PRESSURE MOLAR FLOW COMPOSITION (molar %)
MEK 2-but
79,4°C 100 KPa 155,5 Kmol/h 99,51 154,7567 kmol/h 0,49 0,7565 kmol/h
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4. EQUIPMENT SIZING The sizing of equipment plays a crucial part in any preliminary design procedure. It is important to appropriately size all the individual unit operations so as to aid in the equipment costing in the later stages of design. The sizing of the extraction distillation columns would be explained in detail.
4.1 Extraction column sizing The first step would be determine the continuous and dispersed phase in the column. This is achieved using the Selker and Sleicher correlation (prediction which is based on the phase volume ratio, density and viscosity of each phase). The requisite properties are extracted from HYSYS. (
)
Where are the volume of the light phase and the volume of the heavy phase in consistent units. The light and heavy phases are chosen based on the mass density of each phase. All the calculations have beed clearly demonstrated in the Appendix 3. The results from the calculation shows that X>3.3 which would mean that the heavy phase would be the dispersed phase. Furthermore, the cross sectional area of the extraction column was also calculated. This calculation was done using the superficial velocities for continuous and dispersed phase at flooding condition. The equations presented below were used for the calculations.
[
] (
[
] )
A is first calculated based on the properties and the chosen packing material. Using A, B is estimated. The volume is determined based on the calculated value of B. Using the volumetric flow rate, the cross sectional area of the column is estimated to be 4,2 ft2 which gives a column diameter of about 1,067 m. The height of the column is determined using the following expression. Where N is the number of stages and HETS is the Height Equivalent to a Theoretical Stage.
HETS is calculated using the extraction factor, λ and HTU is 3.5.
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With 10 number of stages, the column height is calculated to be 9,2 m(inclusive of a 2 m allowance for auxillary equipment). For further detailes see appendix 2
4.2 Absorption and Distillation sizing The sizing is done using HYSYS. An extra height of 1,5 m was added for auxiliary equipment. Table 1. Size of absorption and distillation columns.
Column
Diameter(m) Height(m)
Absorption
0,9144
8,127
Distillation 1
0,9144
22,84
Distillation 2
1,829
15,52
Flash vessel sizing A vertical flash separor was sized next. The minimum allowable flash vessel diameter is given as: √ The
settling
velocity
is
estimated
using
the
equation
below
-
√ A demister pad is not used in the design as it is assumed that all particles are vaporized completely in the flash column. Thus the minimum diameter of the vessel vessel without a demister pad is D v=2,042 m. 1 The liquid level is assumed at 0.62 m so the final height of the flash vessel is 2,66 m
Section 10.9.1 & 2 in Sinnot R, G: Coulson & Richardson´s Chemical Engineering series: “Chemical Engineering Design”, 5th ed, Buttherworth-Heineman 2009. 1
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5. Energy Analysis The proposed plant design was then evaluated completely in terms of the energy demand required for running the plant at full capacity. Energy analysis is a very critical juncture in every design procedure, as it can help in identifying options for improving the energy efficiency of the proposed plant layout. Pinch analysis is the most widely used tool to evaluate the energy demands of a complete system. The pinch analysis presents a simple methodology for systematically assessing the chemical process systems and surrounding utilities with the help of the First and Second laws of Thermodynamics. The power of pinch analysis mainly stands in its ability to offer a holistic analysis of the possible heat exchanges in a large and completely integrated system.
Figure 1. Key steps in Pinch Technology
(Figure courtesy of B.linhoff et.al 1983) Using Pinch Technology, it is possible to identify appropriate changes in the core process conditions that can have an impact on energy savings. Targets can also be set for the utility loads at various levels (e.g. steam and refrigeration levels). The utility levels supplied to the process may be a part of a centralized site-wide utility system (e.g. site steam system Pinch Technology therefore provides a consistent methodology for energy saving, from the basic heat and material balance to the total site utility system.
5.1 Base Case The first step in a detailed pinch analysis is to obtain the relevant energy flows in a system. The thermal streams in the process were identified. The three major thermal streams noted in the process were –
Convective Heat Streams Phase Change Streams Reactive Thermal Streams
The heat fluxes for all the thermal streams are summarized and represented in the table below. Table 1. Thermal Streams
Sl no
Stream Tstart Ttarget Q (KW) Butanol Feed 1 Butanol Feed Pre-heating 30 126,6 997,6 2 Butanol feed Evaporation 126,6 126,7 1784 3 Butanol Feed Superheating 126,6 450 2687 Reactor 4 Reactor Stream* 405 450 2092 Two Stage Condensation Of Reactor effluent 5 Condensation Stage 1 405 64 2771 6 Condensation Stage 2 64 33,1 1524 Extractor 7 Raffinate Phase cooling 34,69 24,44 222,7 Distillation Column 1 8 Condenser 79,9 79,6 664,4 9 Reboiler 113,4 114 796,4 10 TCE effluent cooiling 113,4 35 71,45 Distillation Column 2 11 Condenser 79,4 79 3447 12 Reboiler 99 99,5 3748 *The reactor temperature profile is presented in the reverse order to indicate that there requirement needed here.
Type Cold Cold Cold Cold
Hot Hot Hot Hot
Cold Hot Hot Cold is a net heat
The stream data above represents those streams in the process where there is a net heat flux noted. It is assumed that the absorption, extraction and flashing operations are black boxes where there are no noticeable heat fluxes. The stream data presented above is used to construct the Grand composite curve (GCC) for the entire process. The GCC represents the entire process as heat sources and heat sinks. The GCC for the MEK production process is depicted below. Grand Composite Curve
500
Interval Temperature (°C)
450 400 350 300 250
Legend
200
Heat Sink
150
Heat Source
100 50 0 0
2000
4000
6000
Heat Load (KW)(KW) 8000
10000
Figure 2. Grand Composite Curve of the process
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From the above GCC the minimum heating and cooling requirement of the process is estimated as – Minimum Heating Requirement (Qhmin) = 9012 KW Minimum Cooling Requirement (Qcmin) = 5613 KW The pinch Temperature is 74, 9 °C As mentioned earlier the pinch analysis gives a holistic view of the entire energy requirement of the MEK production. A global ΔTmin of 10 °C is chosen for all the Pinch analysis procedures described earlier and henceforth. This value is chosen so that there is sufficient tradeoff between energy savings and capital costs incurred. Based on the preliminary GCC the utility requirement was also assessed. The heating and cooling utilities available are listed below.
LP steam at 4.5 bar MP steam at 15 bar Cold water available at 10 °C (maximum 15 °C at the outlet) Hot utility at high temperature (a “virtual utility” at a temperature higher than 600°C ) A Background foreground analysis with the aforementioned utility system yielded the following results. Back/Foreground curves 700
600
Interval Temperature (°C)
500
400
300
200
100
0 0
1000
2000
3000
4000
5000
6000
7000
8000
9000
10000
Heat Load (KW)
Figure 3. Back/Foreground curves
Based on the above assessment it could be concluded that there was no requirement for an MP header in the process. The result of the Background foreground analysis can be summarized as follows – 15
LP header at 147,9 deg (MP header was not required) = 6419 KW Hot utility at 600 deg c = 2593 KW Cooling water utility at 15 deg = 5612 KW
5. 2 Process Integration Assessments 5.2.1 Separation MER The base case represents a scenario where there is poor process integration within the various system utilities, leading to higher heating and cooling demands. In order to minimize these demands a process integration assessment is needed. Such an assessment was first performed between the 2 Separation units in the process. A Background foreground analysis between the separation units would give an estimate of the Separation MER (Minimum energy Requirement). This analysis yielded the following results –
Legend Distillation Column 1 Distillation Column 2
Figure 5.1. Background Foreground analysis
Interval Temperature (°C)
Complete process integration can be achieved between the distillation columns if the distillation column 2 can be shifted in a such way that the condenser of distillation 2 can provide heat to the reboiler of distillation 1. This temperature shift can be bought about by increasing the pressure drop within the distillation column 2. The completely integrated Separation process is represented below. 180 160 140 120 100 80 60 40 20 0 0
500
1000
1500
2000
2500
3000
3500
4000
Heat Load (KW) Figure 5.2. Background Foreground analyis
The results of the Seperation MER studies is summarized below – 16
The second column temperature was shifted by 50 °C. Complete Process integration was achieved within the separation units. Qhmin = 3748 KW Qcmin = 3315 KW Heat savings = 796 KW 5.2.1.1 Complete Process Integration (Distillation column 2 pressure shifted) A complete Background foreground analysis between the completely integrated separation units and the remainder of the process is done. This is done in order to investigate potential for complete process integration within the process. The foreground would consist of the two separating units (Distillation columns 1 and 2 – Reboiler and condensers) and the background includes the remaining process. The results of this study are presented below. Back/Foreground curves 500 450 400
Interval Temperature (°C)
350 300 250 200 150 100 50 0 0
1000
2000
3000
4000
5000
6000
7000
Heat Load (KW) Figure 5.3. Background Foreground curves
Process integration was achieved by using a foreground shift. As is evident from the figure above, complete integration has been achieved. This was obtained by shifting the foreground. However, this does not accurately quantify the heat savings and the corresponding heating and cooling MER. This is due to the fact that the heat loads in the separation units would be altered when pressure shifts are introduced. Moreover, the number of trays and the external reflux ratio would also be impacted when the pressure is changed leading to different heat fluxes throughout. Thus a more rigorous assessment of the process integration is needed. The results of this preliminary assessment is presented below – Heat savings within separation = 796 KW Heat savings within process = 1748 KW
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Total savings = 2544 KW Qhmin = 6347,38 KW Qcmin = 2947,85 KW Process pinch = 154,4 °C 5.2.1.2 Rigorous Process Integration (Distillation column 2 pressure shifted) A real time simulation of the effects of pressure shift within the separation units is needed in order to accurately quantify process integration avenues. These simulations were achieved in Aspen HYSYS and the corresponding heat fluxes for the real time simulation were used in order to present the rigorous process integration assessments. It is important to consider the number of trays and the external reflux ratio when the pressure is changed. The first changes to the proposed plant layout were suggested in the rigorous process integration assessments. Feed preheating for the respective distillation column was implemented. This was done in order to increase the efficiency of separation within the units. The main motivation for bringing about this change would be that the number of trays in the distillation columns would be drastically reduced. The final process was assessed with this feed preheating. The results of this Background foreground analysis is presented below –
500 450
Interval Temperature (°C)
400 350 300 250 200 150 100 50 0 0
1000
2000
3000
4000
5000
6000
7000
8000
9000
Heat Load (KW)
Figure 5.4. Background Foreground analysis
Complete process integration was achieved. The shifted pressures were introduced real time into distillation column 2. The design of the new column yielded about 110 odd stages. Such a design is practically not feasible. Thus it would be meaningless to invest in such a process. Thus an alternative was investigated. (Column 1 was pressure shifted instead)
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5.2.1.3 Complete Process Integration (Distillation column 1 pressure shifted) Complete process integration can also be achieved between the distillation columns if the distillation column 1 is shifted in a such way that the condenser of distillation 1 can provide heat to the reboiler of distillation 2. This temperature shift can be bought about by increasing the pressure drop within the distillation column 1. The completely integrated Separation process is represented below. 200
Interval Temperature
150 100 50 0 0
500
1000
1500
2000
2500
3000
3500
4000
4500
Heat Load (KW) Figure 5.5. Background Foreground analysis
The results of this Seperation MER studies is summarized below – The first column temperature was shifted by 35 °C. Complete Process integration was achieved within the separation units. Qhmin = 3880 KW Qcmin = 3446 KW Heat savings = 664.4 KW Just as in the previous case, a complete Background foreground analysis between the completely integrated separation units and the remainder of the process is done. The results of this study is represented below -
500 450
Interval Temperature (°C)
400 350 300 250 200 150 100 50 0 0
1000
2000
3000
4000
5000
6000
7000
Heat Load (KW) Figure 5.6. Background Foreground analysis
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The results of this preliminary assessment is presented below – Heat savings within separation = 796 KW Heat savings within process = 1748 KW Total savings = 2544 KW Qhmin = 6347,38 KW Qcmin = 2947,85 KW Process pinch = 154,4 °C To quantify these savings in a more accurate manner, rigorous integration methods need to be assessed. 5.2.1.4 Rigorous Process Integration (Distillation column 1 pressure shifted) A real time simulation of the effects of pressure shift within Distillation column 1 was done. These simulations were achieved in Aspen HYSYS and the corresponding heat fluxes for the real time simulation were used in order to present the rigorous process integration assessments. The results of these assessments is presented below -
500 450
Interval Temperature (°C)
400 350 300 250 200 150 100 50 0 0
1000
2000
3000
4000
5000
6000
Heat Load (KW)
Figure 5.7. Background Foreground analysis
The rigorous distillation yielded complete process integration when the foreground (Separation units with feed preheating) was shifted by about 8 units. The pressure in the second column was shifted by 165 KPa (Total pressure within column 265 KPa). This shift in pressure, lead to a new design for the distillation column. In this new design, 22 trays was suggested, which was more practically feasible 20
than the previous rigorous process integration. Although the overall energy savings was slightly lesser in the second alternative (2129 KW as opposed to 2554 KW), the second alternative was more feasible due to a much lower column height. Thus it would be meaningful to further explore and invest in this alternative. The final results of the process integration analysis along with the GCC for the completely integrated process are provided below. Heat savings within separation = 1130,87 KW Heat savings within process = 1781 KW Total savings = 2129,9 KW Qhmin = 8181,7 KW Qcmin =4633 KW Process pinch = 74 °C Grand Composite Curve 500 450
Interval Temperature (°C)
400 350 300 250 200 150 100 50 0 0
1000
2000
3000
4000
5000
6000
7000
8000
9000
Heat Load (KW)
Figure 5.8 Background Foreground analysis
5.2.2 Capital - Energy Trade-offs The best design for an energy efficient heat exchange network will often result in a tradeoff between the equipment and operating costs. This is dependent on the choice of the ΔTmin for the process. The lower the ΔTmin chosen, the lower the energy costs, but conversely the higher the heat exchanger capital costs, as lower temperature driving forces in the network will result in the need for greater area. A large ΔTmin, on the other hand, will mean increased energy costs as there will be less overall heat recovery, but the required capital costs will be less. The chosen global value of 10 deg C could be
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considered for an initial assessment. Further assessment with other values of ΔTmin would also be needed to validate the pinch studies.
5.3 Heat exchanger Network design Heat exchanger network (HEN) design is a key aspect of chemical process design. Typically, 20-30% energy savings, coupled with capital savings, can be realized in state-of-the-art flowsheets by improved HEN design [1]. The task involves the placement of process and utility heat exchangers to heat and cool process streams from specified supply to specified target temperatures. The objective is to minimize the total costs, i.e. capital and operating costs. The energy targets with maximum energy recovery can be evaluated using Euler’s rule. The equation of Euler’s rule is presented below. NminMER= (NH+NC+NU-1)AP + (NH+NC+NU-1)BP here AP is above the pinch and BP is below the pinch. NH, NC, NU stands for the number of hot and cold streams and number of utilities. Nmin,MER = ( 2+5+2 -1)+(7+3+2-1) Nmin,MER = 19 The pinch principle explains how the process must be separated into two regions, above and below the pinch, for network design in order to achieve the energy targets. The understanding of the pinch location is therefore very important for network design. The data employed in the stream representation for the process is shown below – Table 5.1 Data for the pinch analysis
Tstart
Ttarget
Q
°C
°C
kW
30 126,6 126,6 405 405 34,71 113,4 150,6 79,4 48,67 98,99 64 113,4 34,69
126,6 126,7 450 450 64 93,35 113,2 150,8 79,3 80,33 99 33,1 34,97 24,44
997,6 1784 2687 2092 2771 133,9 1131 1216 3506 245,5 3512 1524 108 222,7
Based on the start and target temperatures mentioned above, the fuzzy diagram for the HEN is created. The primary criterion of choosing a particular HEN was the pinch violations. Several network designs 22
were proposed. The corresponding pinch violations for each were estimated. The final network is presented in the subsequent section. Some of the preliminary network designs and the corresponding pinch violation estimations are presented in Appendix 4.
Figure 5.9 Pinch analysis
The heat exchanger network shown above was obtained based on the basic rules for heat exchange above and below the pinch. Above the pinch: FCpcold > FCpHot Above the pinch: FCpcold < FCpHot Based on the heat exchanger network proposed, the following were the heating and cooling demands of the process – Minimum Hot Utility – 8202 KW Minimum Cold Utility – 4658,7 KW The above network was proposed based on the assurance of minimum pinch violations. The pinch violations can be estimated as – (Minimum Hot Utility + Minimum cold Utility) network – (Qhmin+ Qcmin) Pinch Violations = 46 KW These violations are very much low. Thus this design can be accepted as a probable heat exchanger network.
5.3.1 Practical ramifications of the heat exchanger network The most important step in any heat exchanger network is the practical basis behind the design. The network chosen must have sound practical grounds which would mean that it would be feasible to 23
carry out the design in reality. In the network suggested above the heat from the flue gases leaving the reactor would be used to preheat the reactor feed and in evaporation as well. The condenser of first distillation column would exchange heat with the reboiler of the second column. The heat from this condenser would also be used for preheating the feed that enters both the distillation columns. Thus heat exchange proposed would thus require a complex network of ancillaries. The network proposed was a result of rigorous design. The heat exchangers which had an area of less than 5m2 were removed and the final design presented is the corrected design. To obtain a working simulation of these proposed networks would entail the use of HYSYS. However, due to convergence problems encountered in the simulations, the proposed modifications to the plant design have been represented in a process flow diagram. This flow sheet is presented in Appendix 5.
5.4 Design of reactor heating It is common that chemical plants have to handle an off-gas stream which contains a lot of potential pollutants. In the case of the MEK plant the off-gas has a considerable amount of hydrogen but also of MEK (the loss of the process) and of Trichloroethane. The combustion of MEK and Trichloroethane most likely produces several pollutant species. In addition, the combustion of hydrogen in air is also a big source of thermal-NOx. The waste heat recovery from the combustion of hydrogen-rich off gas from the MEK plant is studied only from an energy point of view. In technical practice, steam production for process heating can be done in a furnace (steam boiler) or in heat recovery steam generators (HRSG). High temperature heating is necessary to carry out the catalytic conversion of Butanol into MEK. The size of the reactor for a given conversion specification depends on the heat transfer between the reactive medium and the hot utility stream through the reactor tubes. Accordingly, the choice of the heating medium might influence the reactor size.
The reactor heating is designed as a burner in which the off-gas is burnt (with primary air) and subsequently mixed with additional air (secondary air). The secondary air must be added in order to decrease the off-gases temperature to a value compatible with the reactor catalyst. Steam can be produced in a HRSG downstream by recovering the heat of the flue gas. Two separate modes of reactor heating have been assessed. The heat from the hot exhaust gases is then used in a steam network to produce the MP and LP steam needed for the above process. The combustor was designed keeping in mind the reactor heating utility profile obtained from HYSYS. The layout of the series heating is illustrated below –
5.4.1 Series Heating The heat from the flue gases obtained above is then used to generate LP and MP steam is a steam network. This steam network, comprising of a HRSG would be used to satisfy the inherent process demands. Any excess steam that remains from this system could then be sold. It must be noted that in order to satisfy the reactor heating profile, natural gas needs to be added. The steam network for the corresponding system is shown in the figure below.
24
Figure 5.11 Series Heating
Figure 5.12 Steam network
The above steam network would also account for losses as well. A GCC of the utility system could then be used to study the excess steam that would be available in this steam network.
GCC the Utilities 600
Temperature
500 400 Process
300
MPs
200
Excess
100 0 0
5000
10000
15000
20000
25000
Q (KW)
Figure 5.12Grand Composite Curve
The GCC would show that about 9940 KW of steam is produced in excess. The representation above is a simplified one. The MP and the LP steam loads are coupled together at the MP level (just for the representation) as it is this MP steam which would be sold. This series heating however has one primary flaw which is the excess consumption of natural gas. The effects of this excessive natural gas demand are elaborated in the economic evaluations in the later sections. 25
5.4.2 Parallel Heating – A more efficient heating arrangement in the form of parallel heating was proposed. This can be imagined as resistances in a parallel circuit, for simplicity sake. This design also incorporates the feed preheating of the secondary air. This secondary air preheating (up to about 250 deg c) would ensure that lower amounts of natural gas are needed at the combustor in order to obtain the reactor heating profile. However, with the incorporation of preheating the corresponding production of excess steam would be lowered. Thus, a tradeoff between the two must be established keeping in mind the economic profitability of the whole project. The layout of the parallel heating is represented in the figure below –
Figure 5.13 Layout of the parallel heating
The heat after the combustion is once again used for inherent process LP and MP steam requirements and this steam is once again generated in a HRSG as in the previous case. The GCC of the utilities would provide a clear picture of the excess steam that is produced as a result of this reactor heating module. This GCC is presented below.
GCC of the utilities
Temperature
600 500 400
Process
300
MPs
200
Air preheating
100
Excess
0 0
5000 10000 15000 20000 25000 Q(KW)
Figure 5.14 GCC of utilities
As expected, the excess steam produced in this case is much smaller compared to the previous case. The amount of excess steam produced was 450 KW. Moreover, the corresponding use of natural gas is also considerable lowered. The economic profitability is described in later sections. 26
6. Process Economics The economic evaluation of the proposed preliminary plant design would provide solid grounds for further development of the design. Before the investors agree to spend large amounts of capital on a proposed project, the management must be convinced that the project would provide a sound investment. Thus process economics is as important as the design procedure itself. In this section the various profitability measures such as the Net payback period, Net future worth, Net present value and the rate of return on investment will be assessed in detail. Based on these indicators the feasibility of the proposed design would be assessed.
6.1 Cost estimations The primary goal of the MEK production plant is to make profit, and an estimate of the investment required is needed before the profitability of this project can be evaluated. Cost estimation would begin with the estimation of the capital costs. This would then be followed by operating cost estimates and finally based on these two; the profitability indicators of the design project are estimated.
6.1.1 Capital Cost Estimates The capital costs is made of two primary components
Fixed capital - the total cost of designing, constructing, and installing a plant and the associated modifications needed to prepare the plant site. [2]
Working Capital - The capital that is tied up in maintaining inventories of feeds, products, and spare parts, together with cash on hand and the difference between money owed by costumers (accounts receivable) and money owed to suppliers (accounts payable). [2]
6.1.1.1 Process Equipment Cost Estimations The costs of all the primary and ancillary equipment involved in the production was estimated first. The estimation of the purchased equipment cost was calculated by:
Ce - Purchased equipment cost a, b - constants S - Size parameter n – Exponent (Particular to an equipment type) The cost estimations were based on data from the year 2007. In order to estimate the prices for the year 2012 the Chemical Process Engineering (CPE) -index was used as follows: Cost in year 2012 =
The CPE indices chosen for the above estimations are represented below. 27
Table 6.1 CPE indices
Year CPE-index (US conditions)
2004 444
2007 525
2012 600
6.1.1.1.1 Reactor The reactor cost is estimated based on the costs for a shell and tube heat exchanger including a factor of 20% to account for the added complexities. The sizing parameter was the heat exchange area which was calculated using the total length of the reactor, the number of tubes and the diameter of each tube (external area of all the tubes). The sizing parameter, S, was the reactor area and was calculated from the following equation:
A -Area, D - Diameter, l – length, nt - Number of tubes Table 6.2 sizing parameters
a Reactor
b
24000
S 46
n
0.9623
length 1.2
1.413
Volume C ($) 0.9623
24043.93
Final C (2004)($) 28852.71
6.1.1.1.2 Absorption column A packed column was designed in HYSYS and it is this column that must be costed. The packing chosen was Ceramic Intalox saddles of 1.5 inches each. The vessel and the packing material are costed separately. The sizing parameter, S, for the vessel was the shell mass and this is estimated with the following equation:
m - Shell mass, DA - Vessel diameter, LA - Vessel length, tw - Wall thickness (including corrosion allowance). Here tw is estimated from –
Where Pi – Internal Pressure (psi), Di – Internal Diameter (mm), E – Joint Efficiency (,85), S – Allowable Stress (ksi) The Sizing parameter for the absorption column is the Section diameter. This value is obtained from HYSYS. The final cost of the column would be the sum of the vessel and packing costs. Due to the presence of highly corrosive species in the column, the material of construction chosen was Stainless Steel 304. Table 6.3 Final cost of the column
Absorption Column Vessel Tray
a
b 10000 110
S 29 380
n 1496.59 1.067
C ($) 0.85 1.8
24495.57 537.0508
Final C (2004)($) 1525.59 381.067
28
6.1.1.1.3 Distillation columns The design proposed has two distillation columns. It is thus necessary to cost both these columns as well. As the case for the absorption column, the tray and the vessel are costed separately. The first column would handle corrosive agents and thus would be constructed in Stainless Steel 304. The other column on the other hand could be constructed from Carbon Steel. Table 6.4 Final cost of the distillation column 1
Distillation Column 1
a
b
S
10000 110
Vessel Tray
29 380
n
C ($)
2574.61 0.9144
0.85 1.8
Final C(2004)($)
32988.08 433.4661
2603.61 380.9144
Table 6.4 Final cost of the distillation column 2
Distillation Column 2 Vessel Tray
a
b
S
10000 110
29 380
n
C ($)
4899.056 1.829
0.85 1.8
Final C(2004)($)
49718.56 1236.596
4928.056 381.829
6.1.1.1.4 Extraction column The extraction column was costed in the same manner as the absorption column. The tray and vessel were costed separately and the final column cost was the sum of these two costs. As is the case above, due to the presence of corrosive species, the material of construction for the column was Stainless Steel 304. Table 6.5 Final cost of the extraction column
Extraction Column
a
b 10000 0
Vessel Packing - Ceramic intalox
S 29 1800
n 793.9 0.7
C ($) 0.85 1
18456.66 1260
Final C(2004)($) 822.9 1800.7
6.1.1.1.5 Flash Separator The Flash separator is costed in the same manner as the absorption column. However only the vessel is costed as there is no packing in such a unit operation. Stainless Steel 304 is used as the material of construction as multicomponent gas and liquid phase change occurs within. Table 6.6 Final cost of the Flash vessel
Flash Vessel Vessel
a
b 10000
S 29
n 362.5
C ($) 0.85
Final C(2004)($) 14343.2 391.5
6.1.1.1.6 Pumps Pumps are needed for transporting and elevating liquid streams (for example usually 2-3 around a column, some of them possibly shared by two columns)In the plant layout proposed in HYSYS all the pumps have not been accounted for, thus an estimate of the pump requirement is first made. About 23 pumps were estimated as an initial guess. This number would also include pumps required for contingency situations such as pump failure. The sizing parameter needed for pump costing was the flow rate. A typical flow (by looking at the different liquid streams) was chosen to estimate the cost.
29
Table 6.7 Final cost of the pump
a
b
S
6900
Pump
n
206
3.611
C ($) 0.9
Final C(2004)($) 7554.232 173747.3
6.1.1.1.7 Fans In the design proposed, the primary requirement of the fans would be in the Combustor (burning the H2 rich flue gas to heat the reactor). Fans would be needed in the primary air and secondary air supplies. The fans were costed based on the air flow rate alone. The value of the flow rates was obtained from HYSYS. The charts given in Plant design and economics for chemical engineers, Max S. Peters and Klaus D. Timmerhaus.4th ed. (McGraw-Hill chemical engineering series) were used for the costing. Centrifugal radial fans with electric drives were chosen. The obtained cost for the fan was about 5500 $(2004). 6.1.1.1.8 Burner/Combustor The cost of burner was modeled as twice the cost of the fan that would be needed at the inlet of the combustor. This would be most logical as the fan cost would be the primary contributor to the burner cost. The cost of the burner would be about 9000 $(2004). 6.1.1.1.9 Heat exchangers The Heat exchangers were costed using the heat exchange area as the Sizing parameter. This costing would be the same as the reactor case. The parameters for the heat exchanger costing are represented below. Table 6.8 Parameters for the heat exchanger
a
b 10000
n 88
1
The final heat exchanger cost is estimated to be about 179979 $(2004). 6.1.1.1.10 Storage Tanks The economic estimation of the tank costs were done based primarily on the feed and the product tanks. This would be a preliminary assessment as other tanks within the process, such as say the intermediate mixing tanks, reserve tanks etc. have not been considered. The sizing parameter for the economic estimations is the capacity. In this case the capacity needed during 24 hours of production is considered. The tanks considered had cone shaped heads and were constructed with Carbon Steel. Table 6.9 Final cost of the distillation column 1
a Butanol feed tank TCE feed Tank MEK product Tank
b 5000 5000 5000
S 1400 1400 1400
n 343.2 10 332.64
C ($) 0.7 0.7 0.7
Final C ($) 88370.26 1743.2 12016.62 1410 86566.19 1732.64
30
6.1.1.2 Purchased Equipment CostThe total purchased equipment cost would be the summation of each individual process equipment cost. Each individual cost is expressed as cost in the year 2012 using the CPE indices. Finally the current exchange rate ($ to SEK – 6,7) is taken into consideration to represent the final cost is SEK. The final purchased cost estimates are represented in the table below. Table 6.10 Final cost of the distillation column 1
Utility
Final C Cost 2012, US$ Cost 2012, ($)(2004) SEK 28852.6318 32974.43634 220928.7235 25032.6246 28608.71383 191678.3827
Reactor Absorption tower- Vessel Absorption tower- Tray Dist1 - Vessel Dist1 - Tray Dist2 - Vessel Dist2 - Tray Extraction Vessel Extraction Packing - Ceramic intalox Flash vessel Pump Furnace Butanol feed tank TCE feed Tank MEK product Tank Fan + Burner Fan Total HX cost TOTAL EQUIPMENT COST
33421.5478
38196.05464
255913.5661
50152.0304
57316.60612
384021.261
19716.6612
22533.32704
150973.2911
391.5 173747.342 95021.89 1743.2 1410 1732.64 14494.9084
447.4285714 198568.391 108596.4457 1992.228571 1611.428571 1980.16 16565.60965
2997.771429 1330408.22 727596.1863 13347.93143 10796.57143 13267.072 110989.5846
179979.6 625696.576
205690.9714 715081.8015
1378129.509 4791048.07
The above consolidated cost estimation is indicative of the contribution of individual unit operations to the total purchased equipment costs. The heat exchangers account for a major portion of the total costs. The furnace is also a primary contributor to the total purchased equipment costs. The individual cost contributions are represented in the figure below. Total equipment cost contribution
MEK product Tank
Total HX cost Fan
TCE feed Tank Butanol feed tank
Furnace
Reactor
Absorption tower
Dist1 Dist2 Extraction Flash vessel Pump
Figure 6.1 Total cost contribution
31
6.1.1.3 Fixed Capital Cost estimates The fixed capital is estimated using the detail factorial method. All the estimates are solely based on the total delivered equipment costs. The fixed capital is in turn divided into two sub classes. These are
The direct-cost - that are incurred in the construction of a plant, such as - Equipment erection, including foundations and minor structural work , Piping, including insulation etc. [2]
The Indirect-cost - that are incurred in the design and engineering, contingency measures. [2]
The detailed factorial cost estimates for the economic estimations are detailed in the table below. Table 6.11 Detailed factorial costs
Item Direct Costs Equipment Delivered Cost Equipment Erection Piping Instrumentation and control Electrical Structures and buildings Lagging and paint Utilities Offsites Site Preperation Total capital cost of Installed Equipment Indirect Costs Design,engineering and construction Contingency Total Fixed Capital Cost -F.C
Factor 1 0.4 0.7 0.2 0.2 0.2 0.1 0.5 0.2 0.1 3.6 1 0.4 5
Thus the total fixed capital cost would be 5 times the delivered equipment cost. This is about 24 Million SEK. 6.1.1.4 Working Capital estimates The working capital is estimated as 15% of the total fixed capital. This would amount to about 3,35 Million SEK. 6.1.1.5 Total Capital Estimates The total capital would be the sum of the fixed and working capital investments. Total Capital Investment = Fixed Capital + Working Capital The total capital investment required is estimated to be about 27 Million SEK.
6.1.2 Operating cost estimates Estimation of the operating costs of production is a key step in determining the profit- ability of a process. The operating costs are in turn divided into two sub classes – Variable Operating Costs - Variable costs of production (operation) are costs that are proportional to the plant output or operation rate. These include the costs of: Raw materials consumed by the process, 32
Utilities—fuel burned in process heaters, steam, cooling water, electricity, raw water, instrument air, nitrogen, and other services brought in from elsewhere on the site, Packaging and shipping—drums, bags, tankers, freight charges, etc.[2] Fixed Operating Costs - Fixed production (operation) costs are costs that are incurred regardless of the plant operation rate or output. Fixed costs include: Operating labor, Supervision—usually taken as 25% of operating labor, direct salary overhead—usually 40% to 60% of operating labor plus supervision. , Maintenance, which includes both materials and labor, and is typically estimated as 3% to 5% of ISBL investment, depending on the expected plant reliability, Property taxes and insurance— typically 1% to 2% of ISBL fixed capital, General plant overhead—charges to cover corporate overhead functions such as human resources, research and development (R&D), information technology, finance, etc. [2] 6.1.2.1 Variable Operating Cost Estimates The raw materials for the process were costed taking into account their flow rate. For instance the price of 2-butanol, process water, TCA and cooling water etc. were estimated using the flow rate, either in kg/h of m3/h (depending on the basis chosen by the management). The price of natural gas and electricity were respectively calculated based on their demand in MW. The catalyst used in the reactor consists of brass pellets. It is decided that 1/3rd of the catalyst will be changed every year. Thus the catalyst is costed accordingly. The natural gas demand is dictated by the combustor design proposed. As mentioned in the earlier sections, two configurations of combustor design were proposed – Series and Parallel heating. Based on the configuration chosen, the demand of natural gas would also be different. A detailed economic evaluation of both these alternatives would be done in the subsequent sections. The price of Natural gas was given in SEK/MW and the Lower heating value of natural gas was obtained from HYSYS (about 50 MJ/kg). In addition, there was also production of excess steam. This was after all the requisite process heat was satisfied. This surplus of steam would be sold at 529 SEK/MWh and would account for sizeable revenues. The electricity consumption was based solely on the pump work. The electricity costs are however considered to be negligible when compared to the overall production costs and have this been ignored in this assessment. The different configurations chosen for the reactor heating were assessed based on economic merit. The series heating case is discussed first. Table 6.12 MEK production aspects for the series heating case
Component 2-butanol Process water Trichloroethane MEK Natural gas requirement (LHV methane) Cooling water requirement Catalyst, m3 renewal each year Steam Production
Flow, m3/h 14.3 0.101 0.06151 13.86 44.2345
Flow, kg/h
919.62 0.27633 9941.4
m3/h m3 KW
11640 100.8 88.58 11220 MW
33
Table 6.13 Detailed variable cost estimations
Raw materials
Price
Unit
Cost
Unit
Cost in SEK/y
2-butanol Process water Trichloroethane Catalyst Utilities
7.6 23 0.94 6600
SEK/kg SEK/m3 US$/kg US$/m3
88464 2.32 83 1823.778
SEK/h SEK/h US$/h US$/y
707712000 18560 4448800 12219.3126 0
Natural Gas
450
SEK/MWh
19905.525
SEK/h
159244200
529.41 SEK/MWh
5263076.6 367.848
SEK/h
-42104612.59
SEK/h
2942784
Excess Steam (REVENUE)
0.4
Cooling water
SEK/m3
Total Variable Operating Cost
832273950.7
6.1.2.2 Fixed Operating Cost estimates The fixed operating costs are estimated next. Table 6.14 estimated fixed operating costs
Labour Number of staff Wages, SEK/month Operating Labor
12 27500 330000
Factor for allowances /overhead Lab cost
165000 70950
Supervision Capital Charges Insurance Royalties Maintanence Taxes
66000 2395524,035 239552,4035 239552,4035 1796643,026 479104,807
Subtotal fixed costs
5782326,675
Direct Production Cost
838056277,4
R&D, Sales overheads
0
838056277,4 Annual Production Costs (SEK/y) It must be noted that there are no R&D costs included. This would primarily be due to the fact that the MEK has a fixed consumer base and market and thus would not need much operation research. Moreover, as the plant proposed is a subsidiary to the mother plant, it is assumed that the sales and other overheads would be borne by them. The annual operating costs would be about 83,8 Million SEK.
34
Table 6.15 Final assessment of the operating cost
Total operating cost, SEK/y
838056277,4
Total operating cost,SEK/kg MEK
9,336634106
6.1.2.3 Economic Evaluation of the Project The cumulative cash flow for the plant during its entire lifetime was assessed next. It was assumed that during the 2 year construction period there would be no production at all, and after these two years the factory was supposed to start operating at a medium capacity of 70% of full production In the next year the production was supposed to be increased to 80% of full capacity, and finally in the fourth the plant would be working at full capacity. The total investement of the plant was spread over the first three years in the following manner as shown in the table below. Table 6.16 total investment of the plant
Investment (SEK)
Year
10923589,6 13654487 2730897,4
2013 2014 2015
For all the economic calculations henceforth, the discount rate was considered to be 10% and all costs and profits were increased by 2% each year. The predicted price of MEK for the production was fixed as depicted in the table below. Table 6.17 Predicted price of MEK for the production
Production, tonne/y
80000-120000 10
Price of MEK, SEK/kg
The Net present worth and Net Future worth (explained in the later section) are plotted as follows-
Profitability Assessment for Series heating scenario 1,00E+08 5,00E+07 0,00E+00
SEK
-5,00E+07
0
1
2
3
4
5
6
7
8
9
10
11
12
-1,00E+08
Net Future Worth
-1,50E+08
Net Present Value
-2,00E+08 -2,50E+08 -3,00E+08 -3,50E+08
Years
Figure 6.2 Profitability assessment for series heating scenario
35
The economic indicators chosen for the profitability assessment are detailed below Payback period: refers to the period of time required for the return on an investment to "repay" the sum of the original investment. The time value of money is not taken into account. It intuitively measures how long something takes to "pay for itself." [2] The payback time in our suggested production was about 7 years. Profit margin: It is the difference between selling price of MEK/Kg and the cost price of MEK/Kg. A low profit margin indicates a low margin of safety: higher risk that a decline in sales will erase profits and result in a net loss, or a negative margin. Profit margin is an indicator of a company's pricing strategies and how well it controls its costs. [2] The profit margin obtained for this case is ,67 SEK/MEK, which is a pretty slim profit margin. This is reflected in the long payback period. NFW: (Net Future Worth) It is the value of an asset at a specific date. It measures the nominal future sum of money that a given sum of money is "worth" at a specified time in the future assuming a certain interest rate. [2] A net future worth of 43 Million SEK is noted in the project. NPW: (net present worth) is defined as the sum of the present values of the individual cash flows of the same entity. A slightly negative NPW is noted. This would in turn indicate that that the given investment would not mature in time to give a profitable investment. Table 6.18 Profitability assessment
Measure of Profitability Pay back time Rate of return Net Future Worth Net Present Value
Value 6,99 14,20% 42 MSEK -80 MSEK
A clear assessment of the table above would indicate that, the proposed plant design with series heating would be a poor choice of investment. This would primarily be due to the fact that the operating costs are very high particularly due to the larger demand for natural gas. The parallel heating scenario will be examined next. The capital cost estimates would be essentially the same for the series and parallel heating options. It is the operating costs that would differ radically. The corresponding costs for the series and parallel heating scenarios are represented in the table below. Table 6.19 Costs for the series and parallel heating scenarios
Series Heating
Parallel Heating 27308974 783210842,2 17,83
27308974 Capital Cost (SEK) 838056277,4 Operating Cost(SEK) 44,2345 Natural gas consumption (MW) A visual inspection of the table would clearly indicate that the profit generated in this second case must definitely be higher. This would be due to the reduced natural gas consumption. The profitability of the two scenarios is presented below.
36
Table 6.20 Profitability of the two scenarios
Measure of Profitability Payback time Rate of return Net Future Worth Net Present Value
Series Heating 6,99 ,14 42 MSEK -80 MSEK
Parallell Heating ,47 years 2,14 640 MSEK 267 MSEK
The shorter payback period and higher rate of return indicate that the second investment would be much more profitable than the first one. Thus the preliminary plant design proposed would include the parallel heating. The Cash flow diagram of this proposed project is described in the figure below –
Profitability Assessment for Paralell heating scenario 7,00E+08 6,00E+08 5,00E+08 4,00E+08 SEK
3,00E+08 2,00E+08
Net Future worth
1,00E+08
Net Present Worth
0,00E+00 -1,00E+08 0
1
2
3
4
5
6
7
8
9
10 11 12
-2,00E+08 -3,00E+08
Years
Figure 6.3 Profitability assessments for the parallel heating scenario
The positive nature of the cash flow indicated in the figure above is a very good trend. This would indicate that the project will constantly generate sizable revenue, and over a span of ten years would give a huge profit. Thus the proposed plant layout can be investigated further in detail and finalized. Tables detailing the complete NFW and NPW are presented in the Appendix 2.
37
7. HIGH PRESSURE STEAM GENERATION A notable feature of the design proposed is the production of excess steam. As illustrated earlier, this steam can be sold to consumers outside the plant battery limits in order to generate sizable revenue. There is an alternative to selling steam. Electricity could be sold instead (0,65 SEK/KWh), and this could generate higher revenues when compared to steam. It is this scenario that would be investigated here. The key objective of this assessment would be:
To investigate the production of superheated steam at high pressure (60 bars, 500°C) to be used in a backpressure steam turbine.
Evaluate the net electricity production.
Possibility of increasing electricity production by buying natural gas.
Estimate the effects on process economics.
Each of the above objectives would be discussed in detail.
7.1 Methodology The high pressure steam would be generated using the heat from the combustor module in parallel heating. The heat which was used to produce MP and LP steam, would instead be used to produce high pressure steam. This high pressure steam would then be sequentially expanded in a steam turbine to generate electricity. The steam network that would bring about this is represented in the figure below.
Figure 7.1Steam network
We work with a stream of water that, after its pressure has been increased up to 60 bar with a pump, is preheated until its saturation point. Next this stream is evaporated, and finally it is superheated to 500 C. The steam then goes across a backpressure turbine in order to generate electricity. The respective MP and LP steam draw off have been expressed in the figure above. The potential to optimize the above system in order to obtain maximum electricity will be evaluated in detail in the forthcoming section.
38
7.1.1 Base Case – High Pressure Steam Generation: Parallel Heating A background-foreground assessment is performed in order to estimate the total potential for the electricity production. The combustor module as chosen as the background and the steam network (shown above) was chosen as the foreground. The GCC of this case is represented below –
Figure 7.2High Pressure Steam Generation: parallel heating
The above figure indicates that there is a scope for optimization in the above process. The background can provide such a large amount of heat which is not being utilized to the maximum potential. Thus, in order to have maximum power production, the foreground must pinch with the background. This can be achieved by increasing the flow rate of the steam in the steam network. In this manner, by changing the steam flow rate, the corresponding electricity production was also varied. The base case optimized to maximum electricity production is represented below.
Figure 7.3 Maximum electricity production
The electricity produced in the above case is about 1182 KW. The capability to increase electricity production is investigated next. The primary manner in which this can be done would be by reducing the secondary air preheating. This would in turn increase the amount of natural gas fed to the combustor. Thus higher electricity production can be done at the cost of increased natural gas consumption. Two cases have been investigated.
Natural gas consumption increased by 25 %
Natural gas consumption increased by 50%
A similar background-foreground analysis as for the base case is performed yielding the following result. 39
Figure 7.4 Trials in Background/Foreground curves
7.1.2 Practical Ramifications of the background-foreground analysis This analysis clearly represents the fact that the MEK process can provide a large amount of heat to the steam network. The practical realization of this analysis would be that the LP and the MP steam demand have been satisfied by the steam produced in the network. There must be a heat exchange between the combustor flue gases and the evaporation and secondary air preheating.
7.2 Economic Evaluation – 7.2.1 Electricity vs. Steam The Investigations detailed above were then assessed around an economic framework. The primary assessment made was to establish whether it would be more profitable to sell electricity as opposed to steam. The base case scenario was compared with the MP steam generation. This profitability assessment was based on the revenues and profit margins from steam and electricity respectively.
Figure 7.5 Revenues
Clearly, the revenues from electricity are much higher than the corresponding revenues from steam. Moreover, the profit margin for the electricity case is also slightly higher. The slight increase in margin (although revenues from electricity are several magnitudes higher) is due to the fact that the capital cost would increase drastically due to the high cost of the steam turbines. Nevertheless, selling electricity would still be a better option when compared to selling steam.
40
7.2.2 Augmenting electricity production As discussed in the methodology, the capability to increase electricity production by buying more natural gas was also examined. An economic assessment of this scenario would present the basis as to whether it is smart in doing such an expansion. Three metrics were studied based on which the final judgment was made. These were – 7.2.2.1 Comparison of Natural gas flow and Total annual production costs The different trials were compared based on the total production costs. This would provide a metric as to how the natural gas costs affect the overall production costs. The following results were obtained-
Figure 7.6 Production costs
The above results indicate that there is a sharp increase in the annual production costs when the natural gas flow is increased. 7.2.2.2 Comparison between Natural gas costs and revenues from electricity This metric would give the relative magnitude of the two costs. If the magnitude of the revenues from electricity would be much smaller than the increment in the production costs, as is the case here, it would be meaningless to invest in increasing the capacity of electricity production.
Figure 7.7comparison between Natural costs and revenues from electricity
41
Profit margin vs. Electricity production This was the final metric that was chosen. This would be a very clear indicator of the economic effects of increasing the electricity production capacity. The results of this metric are presented below –
Figure 7.8 final metric
The profit margin sharply declines with increase in electricity production. Thus it would not be a wise choice to invest in augmenting electricity production. A cumulative decision made based on the three chosen metrics would be that, it would be wasteful to invest in increasing the electricity production. This would primarily be due to the increased natural gas consumption, which would lead to higher operation costs and also due to the increased capital costs accounting for the turbines.
42
9 Conclusions The proposed plant design can be considered to be a very profitable investment. The design would indicate that, the required production of 90000 tons/year at a reactor conversion of 93% is met with ease. The energy analysis on the other hand, is also indicative of very budding trends. There have been significant energy savings obtained with the process integration methods proposed. The heating media for the reactor were also chosen purely on the basis of the energy analysis. The parallel heating module would be recommended due to the lower consumption of natural gas in this configuration. This integration coupled with the design of reactor heating also leads to strong economic incentives with incremented profit margins. The proposed plan to sell steam or electricity could be a debatable aspect. On the one hand the revenues from electricity are much higher than revenues from steam, the capital investment involved in the electricity case would substantially be higher. As indicated in the studies, the profit margin/Kg MEK is not very different. Nevertheless, it would be slightly better to sell electricity as opposed to steam.
43
10 References 1) B. Linnhoff (1983), The Pinch design method for Heat exchanger networks, Chemical engineering Science Vol.38, No.5, pp. 745-763 2) Sinnot R, Towler, G: Coulson & Richardson’s Chemical Engineering series: “Chemical Engineering Design”, 5th ed., Butterworth-Heineman 2009., Second Edition. DOI: 10.1016/B978-0-08-096659-5.00001-8
44
11. APPENDIX 11.1 APPENDIX 1: MASS BALANCE The following conditions are assumed (Group 9):
Production of methyl ethyl ketone, MEK: 90000 tons/year Conversion in the reactor: 93 % Raw material: 2-butanol Product specification, produced MEK: 99 wt% MEK Suggested product specification, recycled 2-Butanol: 99 mol% 2-butanol Loss of MEK from the plant: 1 wt% based on feed Operating time: 8 000 h/year
4
1
2
REACTOR
3
6
SEPARATION
5
Molecular weights: 2-butanol 74 g/mol MEK 72 g/ mol
Reaction:
2-but
MEK
H2
We assume that this is the only reaction. With this data, mass balance can be calculated.
45
MASS BALANCE: First of all, we will calculate the product stream: Product Stream (6) (99 wt% MEK, 1 wt% 2-But) o o
11,1375 tons/h MEK = 11137,5 kg/h MEK 0,1125 tons/h 2-but = 112,5 kg/h 2-but
MEK produced (product stream + losses) If we look at the reaction:
MEK
2-but
H2
F2
-
-
xF2
xF2
xF2
(1-x)F2
156,234
o
2-butanol feed to the reactor
o
2-butanol without react
156,234
2-but 2-but = 870,17 kg/h 2-but
With this, mass balance of 2-butanol can be planted in the separation unit:
There is no 2-butanol in stream 4, and we have how much 2-butanol there is in stream 3 (2-butanol without react) and stream 6 (1 wt% 2-butanol in the product) so: (99 mol% 2-but in stream 5) ;
Now we can plant mass balance of 2-butanol at the beginning of the process:
With this, all streams can be calculated planting mass balance:
46
So, finally we have:
STREAM (Kg/h) 2(Kmol/h) BUTANOL (Kg/h) MEK (Kmol/h) (Kg/h) H2 (Kmol/h) (Kg/h) (Kmol/h) TOTAL (Kmol/year)
1 11674,24 157,76 11674,24 157,76 1262080
2 12432 168 7,445 0,1034 12439,45 168,1034 1344827,2
3 870,24 11,76 11248,85 156,234 312,47 156,234 12431,56 324,228 2593824
4 111,38 1,547 312,47 156,234 423,85 157,781 1262248
5 757,76 10,24 7,445 0,1034 765,205 10,3434 82,75
6 112,5 1,52 11137,5 154,7 11250 156,22 1249760
47
11.2 APPENDIX 2. Economics The estimation of the purchased equipment cost was calculated by: Ce=a+bSn
(1)
In order to calculate the design, size parameter S, the thickness had to be calculated for the following units: absorption tower, distillation column 1 and 2, extractor and flash. The following equation was used: (2)2
11.2.1 Equipment: 11.2.1.1 Absorption tower Table 11.1: material properties for the absorption tower
material properties: weight (kg) material thinkness (mm)
1496,586 SS 304 7*
The design parameters for the calculation of the total cost of the absorption tower calculated with eq. (2) Table 11.2: final size and total cost of the absorption tower
Process unit:
Absorption tower 13 na
Number of trays: Tray height, m: Extra height/top & bottom, m Total height, m
1,5 8,127
Diameter, m
1,067
11.2.1.2 Distillation column 1 Table 11.3: material properties for the distillation column 1
material properties: weight (kg) material thinkness (mm)
2574,61 SS 304 5*
Sinnot R, G: Coulson & Richardson´s Chemical Engineering series: “Chemical Engineering Design”, 5th ed, Buttherworth-Heineman 2009p. 1005 eq. 13.41 *The mm thickness values were estimated using longitudinal stress and material strength factors 2
48
Table 11.4: final size and total cost of the distillation column 1
Process unit:
Distillation tower 1 35
Number of trays:
na
Tray height, m:
1,5 22,84 0,9144 2574,61
Extra height/top & bottom, m Total height, m Diameter, m Vessel weight, kg
11.2.1.3 Distillation column 2 Table 11.5: design parameters for the calculation of the total mass of the distillation column 2
material properties: weight (kg) material thinkness (mm)
4899,08 Carbon Steel 7*
Table 6: final size and total cost of the distillation column2
Process unit: Number of trays: Tray height, m: Extra height/top & bottom, m Total height, m Diameter, m Vessel weight, kg
Distillation tower 2 23 na 1,5 15,52 1,829 4899,056
11.2.1.4 Extraction column Table 11.7: parameters for the absorption tower.
material properties: weight (kg) material thinkness (mm)
Extraction Vessel 193,9 SS 304 5*
Table 11.8: final size and total cost of the extraction column
Process unit: Packing height, m Diameter, m Extra height/top & bottom, m Total height, m Packing volume, m3 Vessel weight, kg
Extraction na 0,7 1,5 9,2 793,9
49
11.2.1.5 Flash
11.2.1.6 Pumps
Table 11.9: material properties for the flash
material properties: weight (kg) material thinkness (mm)
Table 11.10: material properties for the pumps
material properties: weight (kg) Number of pumps thinkness (mm)
362,5 SS 304 9*
1496,586 carbon steel 7*
It is a centrifugal pump.
Table 11.11: final size and total cost of the pump
Process unit Number of pumps (incl replacement for each pumps) Standard pump capacity, m3/h Head, bar
Pump 23 13 na
Table 11.12: final size and total cost of the fan
Process unit Number of fans Capacity, m3/h
Fan 2 145,6
Burner 1 13,89
The material of both feed tanks and product tank are made of cone roof.
Table 11.13: final size and the total cost of the tanks
Process unit Type
Tanks Butanol feed tank
Tanks TCE Product tank
Number Capacity, m3
1 343,2
1 1,476
Tanks MEK Product tank 1 332,64
11.2.2 Operating cost Table 11.14: parameters for the absorption tower. Vessel and tray
Process conditions Component 2-butanol
Flow, m3/h Flow, kg/h 14,3 11640
50
0,101 100,8 0,06151 88,58 13,86 11220 44,2345 MW
Process water Trichloroethane MEK Natural gas requirement (HHV methane) Cooling water requirement Catalyst, m3 renewal each year Steam Production
Utilities sold
919,62 m3/h 0,27633 m3 9941,4 KW
Unit
Excess steam sold
Year
350 SEK/MWh 2010
Cost
Unit
1740,26313 SEK/h
Cost/Year 2010 13922105,04
Total operating cost, SEK/y
521454602
Total operating cost,SEK/kg MEK
69527,28021
11.2.3 Summary costs and revenues (NPW and NFW): Table 11.15: summary of the costs and revenues (NPW and NFW)
Year Capital investment, SEK Operating costs, SEK Production, kg/y Sales income, SEK/y Net cash flow, SEK/y Cumulative cash flow, SEK Discounted cash flow, SEK/y Cumulative DCF, SEK (Year no) Discount rate
Year Capital investment, SEK Operating costs, SEK Production, kg/y Sales income, SEK/y
2013 10 923 589,60
2014 13 654 487,00
2015 2 730 897,40
2016 0,00
0,00 0,00 0,00
0,00 0,00 0,00
-10 923589,60 -10 923 89,60 -10 923 89,60
-13 654 487,00 -24 578 076,60 -12 413 170,00
838056277,4 62 832 000,00 634 603 200,00 -206 183 975 -230 762 051 -170 399 979
854 817 403 76 296 000,00 770 589 600,00 -84 227 803 -314 989 854 -63 281 595
-10 923589,60 0 0,1
-23 336 759,60 1
-193 736 739 2
-257 018 334 3
2017 0
2018 0
2019 0
2020 0
871 913 751,00 89 760 000,00 897 600 000,00
889 352 026,02 89 760 000,00 920 040 000,00
907 139 066,54 89 760 000,00 943 041 000,00
925 281 847,87 89 760 000,00 966 617 025,00 51
Net cash flow, SEK/y Cumulative cash flow, SEK Discounted cash flow, SEK/y Cumulative DCF, SEK (Year no) Discount rate
25 686 249,00 kr -289 303 605,34
30 687 973,98 -258 615 631,36
35 901 933,46 -222 713 697,91
41 335 177,13 -181 378 520,78
25 686 249,00
19 054 817,40
20 265 705,47
21 211 481,71
-231 332 084,77 4,00 0,10
-212 277 267,37 5,00
-192 011 561,89 6,00
-170 800 080,18 7,00
Year Capital investment, SEK Operating costs, SEK Production, kg/y Sales income, SEK/y Net cash flow, SEK/y Cumulative cash flow, SEK Discounted cash flow, SEK/y Cumulative DCF, SEK (Year no) Discount rate
2021 0
2022 0
2023 0
2024 0
943787484,8 89760000 990782450,6 46994965,79 -134383555
962663234,5 89760000 1015552012 52888777,36 -81494777,63
981916499,2 89760000 1040940812 59024312,97 -22470464,66
1001554829 89760000 1066964332 65409503,29 42939038,63
21923498,38
22430004,52
22756427,78
22925631,87
-148876581,8 8 0,1
-126446577,3 9
-103690149,5 10
-80764517,64 11
Year Capital investment, SEK
2013 10 923 589,60
2014 13 654 487,00
2015 2 730 897,40
2016 0,00
Operating costs, SEK Production, kg/y Sales income, SEK/y
0,00 0,00 0,00
0,00 0,00 0,00
Net cash flow, SEK/y
-10 923 589,60 -10 923 589,60 -10 923 589,60 -10 923 589,60
-13 654 487,00
749552077,4 62 832 000,00 634 603 200,00 -117 679 775
764 543 119 76 296 000,00 770 589 600,00 6 046 481
-24 578 076,60
-142 257 851
-136 211 370
-12 413 170,00
-97 256 012
4 542 811
-23 336 759,60
-120 592 772
-116 049 961
1
2
3
Cumulative cash flow, SEK Discounted cash flow, SEK/y Cumulative DCF, SEK
0 0,1
(Year no) Discount rate
Year Capital investment, SEK Operating costs, SEK Production, kg/y
2017 779833981,3 89760000
2018
2019
795430660,9 811339274,2 89760000 89760000
2020 827566060 89760000 52
(NFW)
(NPW)
Sales income, SEK/y Net cash flow, SEK/y Cumulative cash flow, SEK Discounted cash flow, SEK/y Cumulative DCF, SEK (Year no) Discount rate
Year Capital investment, SEK Operating costs, SEK Production, kg/y Sales income, SEK/y Net cash flow, SEK/y Cumulative cash flow, SEK Discounted cash flow, SEK/y Cumulative DCF, SEK (Year no) Discount rate
897600000 117766018,7 18445351,66 117766018,7
920040000 943041000 124609339,1 131701725,8 106163987,4 237865713,2
966617025 139050965 376916679
1716057,561 4 0,1
79088653,23 5
153430844 6
224785976 7
2021
2022
2023
2024
77372595,67 74342190,78 71355131,7
844117380,8 89760000 990782450,6 146665069,8 523581748,4
860999728,5 878219723 895784117 89760000 89760000 89760000 1015552012 1040940812 1066964332 154552283,4 162721089,2 171180215 678134031,8 840855120,9 1012035336
68420337,47
65545255,32 62736023,97
59997621,1
293206313,2 8 0,1
358751568,5 421487592,5 9 10
481485214 11
NFW
NPW
53
11.3 APPENDIX 3: Sizing 11.3.1. Extraction column sizing Determination of dispersed and continuous phase
(
)
(1)
The specific properties below are found in HYSYS and the phases are aqueous and organic. Table 11.16 specific properties
Top
Bottom
When X>3.3 the heavy phase will always dispersed.
Calculation of cross-sectional area The following expression have been used [ ] (
[ ]
(2) )
(3)
54
Choose of packing The chosen packing material is 1,5 in. Intalox@ Saddles.3
From the figure 16 the value for B can be determined.
(
(
)
)
(
)
Figure 1. The relation between A and B
Table 6.6. in Sinnot R, G: Coulson & Richardson´s Chemical Engineering series: “Chemical Engineering Design”, 5th ed, Buttherworth-Heineman 2009. 3
55
11.3.2Calculation of the extraction column height. The following expression where used:
HETS stands for Height Equivalent to a Theoretical Stage and N for number of stages-
To be able to determined the extraction factor for the whole column, an mean extract and raffinated phase is calculated between the top and bottom.
The HTU should be determined from the figure provided. The phase velocity calculated “would not intersect with the curve, thus an estimed value of HTU of 3.5 ft was chosen.
56
,77 m
The number of stages in the extraction column is 10 which will give an column height of about 9 m.
Inclusive of auxillary height of 1,5 m the final coulmn height is 9,2 m.
11.3.3. Flash vessel sizing The minimum allowable flash vessel diameter is gives as: √ The diameter of the vessel must be large enough to slow the gas down to below the velocity at wich the droplets will settle out.
√
In this project no demister pad is used and
√
57
The holdup time is 10 minutes.
+ extra level
Figure 2 Flash Vessel sizing
58
11.4 APPENDIX 4: Heat Exchanger Network Design Preliminary Designs Network 1 Figure 3. The relation between A and B
Minimum Hot Utility –8817 KW Minimum Cold Utility –5012 KW Pinch Violations = 1015 KW Network 2
Minimum Hot Utility –8512 KW Minimum Cold Utility –4615 KW Pinch Violations = 313 KW 59
60
11.5 APPENDIX 5: Flow sheet
11.6 APPENDIX 6: heat exchanger area estimation
A 37,15 9,18 15,8 154,9 64,99 11,15 14,39 29,32 66,93 77,16 65,07 2,965 19,344
The values of the areas have been calculated with this equation:
The following indicative values of U (W/m2*K) can be considered:
∆T 141,25 42,88 47,09 14,27 44,77 281 171 488 485 66,8 34,4 53,5 16,92
Q 1784 133,9 245,5 751,59 987 2687 2092 1216 2759 3506 1524 108 222,6
