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Feasibility studies KET050 Optimization of formaldehyde plant Company: Haldor Topsøe Advisors: Christian Hulteberg, Ola Wallberg (Chem.Eng), Mads Kaarsholm (Haldor Topsøe)
Jasmin Aspen, Sandra Gao, Yasmin Hawari, Emmy Lam and Tara Larsson 2015-‐06-‐12 LTH
Abstract This project was carried out as a project in feasibility studies. The aim with this project was to optimize a urea-formaldehyde plant at Haldor Topsøe. The optimization was made with respect to energy and heat integration. The project also included an extensive economic analysis examining possible new investments and improvements.
The process was simulated and optimized using software Aspen plus V8.2 and Aspen Energy Analyzer. The energy optimization was performed through different stages: first evaluation of the existing process using pinch analysis through software, and then the suggestions of the new design were presented; at last the new design will be studied in regard to both energy and economics in order to determine profitability. The economic analysis was by made estimating investment cost from public databases and the payback method.
The suggested new heat exchanger network has better energy utilization, where energy in the tail gas has been utilized and air cooled-off energy from the old network has been used to produce steam in the new one. The excess heat would be enough to produce about 13,000 MWh steam per year, which is 4% more than the old system. The new heat exchanger network will cost about 5,189 USD more than the old one. This heat exchanger network was studied with two options for the produced steam, either selling the steam or using it to produce electricity using a steam turbine. Both ideas are affected by the electricity and steam prices. The calculations showed that installing a steam turbine would not be profitable at all; the high investment cost is hard to recover, giving payback times ranging from 10-16 years depending on which pressure of steam that is used. Selling steam would generate an income of approximately 270,000 USD the first year and 400,000 USD remaining years, which is more profitable and the payback time is around 5 month.
Contents Abstract ..................................................................................................................................................
Contents ..................................................................................................................................................
Abbreviations .........................................................................................................................................
1. Introduction ................................................................................................................................. 1 1.1 Aim .................................................................................................................................. 1
1.2 Disposition ...................................................................................................................... 1
2. Background ................................................................................................................................. 2
2.1 Silver catalyst process ..................................................................................................... 2
2.1.1 Water ballast process ............................................................................................ 2
2.1.2 Methanol ballast process ...................................................................................... 2
2.2 The metal oxide catalyst process ..................................................................................... 2
2.2.1 The Topsøe formaldehyde process ....................................................................... 3
2.3 Process differences .......................................................................................................... 3 3. Method ......................................................................................................................................... 5
3.1 Pinch ................................................................................................................................ 5
3.2 Aspen ............................................................................................................................... 6
3.3 Aspen Simulation and modelling .................................................................................... 6
3.3.1 Reactors ................................................................................................................ 6
3.3.2 Absorption ............................................................................................................ 8
3.4 Heat transfer coefficient calculations .............................................................................. 9
3.5 Calculations for steam turbine ....................................................................................... 11 3.6 Calculations for cost estimation .................................................................................... 13
4. Results ........................................................................................................................................ 14
4.1 Energy optimization ...................................................................................................... 14
4.1.1 Evaluation of the existing process ...................................................................... 14
4.1.2 Redesign of the existing heat exchanger network .............................................. 15
4.1.3 Steam and electricity production ........................................................................ 18
4.1.3.1 Steam ........................................................................................................... 18
4.1.3.2 Steam Turbine .................................................................................................. 18 4.1.4 Heat and Mass balance ....................................................................................... 20
4.2 Cost analysis .................................................................................................................. 21
4.2.1 Heat exchanger calculation ................................................................................ 21
4.2.2 Electricity production ......................................................................................... 24
4.2.3 Sensitivity analysis ............................................................................................. 25
4.2.3.1 Electricity production .................................................................................. 25
4.2.3.2 Steam production and selling ...................................................................... 28
5. Discussion ...................................................................................................................................... 30
6. Conclusion ..................................................................................................................................... 32
7. References ..................................................................................................................................... 33
Appendix 1 ........................................................................................................................................... I Appendix 2 ......................................................................................................................................... III
Appendix 3 ......................................................................................................................................... IV
Appendix 4 ......................................................................................................................................... VI
Appendix 5 ....................................................................................................................................... VII
Abbreviations 𝛼 = ℎ𝑒𝑎𝑡 𝑡𝑟𝑎𝑛𝑠𝑓𝑒𝑟 𝑐𝑜𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑡 (𝐻𝑇𝐶) !
!!∙!
𝑑 = 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑝𝑖𝑝𝑒 𝑚
𝜆 = 𝑡ℎ𝑒𝑟𝑚𝑎𝑙 𝑐𝑜𝑛𝑑𝑢𝑐𝑡𝑖𝑣𝑖𝑡𝑦 !!∙!
𝜌 = 𝑑𝑒𝑛𝑠𝑖𝑡𝑦 !"!!
𝑢 = 𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 !!
𝜇 = 𝑡ℎ𝑒𝑟𝑚𝑎𝑙 𝑣𝑖𝑠𝑐𝑜𝑠𝑖𝑡𝑦 !"!∙!
𝐶! = 𝑒𝑓𝑓𝑒𝑐𝑡𝑖𝑣𝑒 ℎ𝑒𝑎𝑡 𝑐𝑎𝑝𝑎𝑐𝑖𝑡𝑦 !!"∙!
𝑝!" = 𝑐𝑟𝑖𝑡𝑖𝑐𝑎𝑙 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒 𝑏𝑎𝑟
𝛼 = ℎ𝑒𝑎𝑡 𝑡𝑟𝑎𝑛𝑠𝑓𝑒𝑟 𝑐𝑜𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑡 !!!∙!
𝜌! = 𝑔𝑎𝑠 𝑑𝑒𝑛𝑠𝑖𝑡𝑦 !!!!
𝜌! = 𝑙𝑖𝑞𝑢𝑖𝑑 𝑑𝑒𝑛𝑠𝑖𝑡𝑦 !"!!
𝜎 = 𝑠𝑢𝑟𝑓𝑎𝑐𝑒 𝑡𝑒𝑛𝑠𝑖𝑜𝑛 !!
𝐿 = 𝑙𝑒𝑛𝑔𝑡ℎ 𝑜𝑓 𝑡𝑢𝑏𝑒 (𝑚)
𝑔 = 𝑔𝑟𝑎𝑣𝑖𝑡𝑎𝑡𝑖𝑜𝑛𝑎𝑙 𝑎𝑐𝑐𝑒𝑙𝑎𝑟𝑎𝑡𝑖𝑜𝑛, 9.81 𝑚𝑠!
𝜂!"#$%&'(!) = 𝑖𝑠𝑒𝑛𝑡𝑟𝑜𝑝𝑖𝑐 𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑦
ℎ𝑖 = 𝐸𝑛𝑡𝑎𝑙𝑝𝑦 (!"!")
𝐺 = 𝐼𝑛𝑖𝑡𝑖𝑎𝑙 𝐼𝑛𝑣𝑒𝑠𝑡𝑚𝑒𝑛𝑡
𝑎! = 𝐴𝑛𝑛𝑢𝑎𝑙 𝑃𝑎𝑦𝑚𝑒𝑛𝑡
𝑖! = 𝐼𝑛𝑡𝑒𝑟𝑛𝑎𝑙 𝑖𝑛𝑡𝑒𝑟𝑒𝑠𝑡
𝐶! = 𝑃𝑢𝑟𝑐ℎ𝑎𝑠𝑒𝑑 𝑒𝑞𝑢𝑖𝑝𝑚𝑒𝑛𝑡 𝑐𝑜𝑠𝑡
𝐹!"! = 𝑀𝑜𝑑𝑢𝑙 𝐹𝑎𝑐𝑡𝑜𝑟
𝐾$,!"#$ = 𝑇ℎ𝑒 𝑡𝑜𝑡𝑎𝑙 𝑐𝑜𝑠𝑡 𝑜𝑓 𝑡ℎ𝑒 𝑝𝑙𝑎𝑛𝑡 𝑡ℎ𝑖𝑠 𝑦𝑒𝑎𝑟
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1. Introduction This project was carried out as the course KET050 Feasibility studies on Industrial Plants in collaboration with the company Haldor Topsøe A/S. The project was performed over a 14 week period. The task was to optimize a formaldehyde plant at Haldor Topsøe with respect to energy and heat integration. With today’s increasing energy usage the importance to use the energy as effectively as possible is also increasing. Energy optimization is beneficial from an environmental and economical point of view. This is important for a sustainable development of a company.
The optimization was based on a flow sheet for urea-formaldehyde production provided from Haldor Topsøe. At the beginning of the project a visit to the company was made were more information about the process and the task was given.
Haldor Topsøe is a Danish catalysis company that was started in 1940 by Haldor Topsøe. Today the company is located in Kgs. Lyngby north of Copenhagen and their main task is to produce heterogeneous catalysts of different kind, where they have up 150 different catalysts and the second task is to design process plants based on catalytic processes for their customers’ need [1].
1.1 Aim The aim of this project is to optimize the Haldor Topsøe formaldehyde plant. The optimization will be performed with respect to energy use in order to improve the production cost and overall performance. The optimization will be done using pinch analysis technology. A heat and mass balance will be performed over the existing system and the energy flows will be analysed and expressed in Sankey diagram. Possible new design options and energy integration improvements will be suggested for the plant in order to improve the performance.
More specific, this report is focus on finding a better heat exchanger network to possibly release high temperature energy for production of steam and electricity.
The profitability of the new design options will be quantified by estimating the costs for the new equipment and potential saving and incomes from steam production and electricity production.
1.2 Disposition The first part of the report describes the background of two kinds of formaldehyde production and differences between them. The second part describes how the study was performed including the energy optimization and cost estimations. In the third part the results are presented. And lastly the given results are discussed and evaluated and references given
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2. Background Formaldehyde can be produced from methanol by two reactions: Methanol dehydrogenation (2.1) and partial oxidation of methanol (2.2) [2].
Dehydrogenation of methanol:
𝐶𝐻!𝑂𝐻 → 𝐻!𝐶𝑂 + 𝐻! ∆𝐻! = +84 𝑘𝐽/𝑚𝑜𝑙 (2.1)
Partial oxidation of methanol:
𝐶𝐻!𝑂𝐻 +!!𝑂! → 𝐻!𝐶𝑂 + 𝐻!𝑂 ∆𝐻! = −159 𝑘𝐽/𝑚𝑜𝑙 (2.2)
This is done primarily by two major processes either using a silver catalyst or a metal oxide catalyst. In the silver process the methanol goes through both dehydrogenation and partial oxidation whereas only partial oxidation occurs in the metal oxide process [2].
2.1 Silver catalyst process In the silver catalyst process the methanol and oxygen passes over a silver catalyst at about 600-720 °C and atmospheric pressure. The reaction is carried out with excess methanol [3]. The formaldehyde is formed both by the partial oxidation according to equation 2.2 (50-60%) and by the dehydrogenation according to equation 2.1. The overall yield of the process is 86-90% [4].
The extent to which the dehydrogenation and partial oxidation occur depends on the process data. The endothermic dehydrogenation reaction is highly temperature-dependent and the conversion increases with increasing temperature. The amount of process air controls the desired reaction temperature and the extent to which the endothermic reactions occur. The silver process can be divided into two different processes; The Water ballast process and methanol ballast process [3].
2.1.1 Water ballast process In the Water Ballast Process vaporized methanol together with air and water steam passes through a bed of silver catalyst crystals which has a temperature between 550-720 °C and atmospheric pressure. The conversion of methanol to formaldehyde takes place when passing through the catalyst bed. The gaseous formaldehyde is cooled before entering the absorption tower. The formaldehyde is absorbed in different types of solutions depending on what type of product that is wanted [2].
2.1.2 Methanol ballast process In the Methanol ballast process there is only methanol and air in the feed stream but the rest of the process is the same. The Water Ballast Process gives higher conversion. The Methanol Ballast Process can be used in countries were the water supply isn’t as big and the higher conversion will not make up for the costs of the water [2].
2.2 The metal oxide catalyst process The partial oxidation reaction takes place at atmospheric pressure and 270-400 °C and it is highly exothermic with a reaction heat of -159 kJ/ mole. The oxidation takes place with excess air in the presence of a metal catalyst. The most important catalysts are iron-molybdenum or vanadium oxide. The catalysts are modified with small amounts of other
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metal oxides such as V2O5, Cr2O3, CuO, P2O5, and CoO [2]. The overall yield of the process is 88-92% [4].
There are many variations of the metal oxide process e.g. the Perstorp Formox process (single reactor design) and the Haldor Topsøe SR formaldehyde. The type of metal oxide processes that will be discussed in this report is the Haldor Topsøe formaldehyde process since that is the process used for this energy optimization.
2.2.1 The Topsøe formaldehyde process The Topsøe formaldehyde process has a series-reactor design. The general process is descripted below and a simplified process scheme is showed in Figure 1.
Reactor Reactor AbsorptionColumn
Chilling Absorber
CatalyticIncinerator
MeOH
TailGas
Formalin
Urea orProcess water
Excess Water
CleanedTail Gas
Air
Excess Air
Figure 1: Simplified flow sheet of the Haldor Topsøe formaldehyde process
Methanol is first evaporated in a gas heater, which is heated by hot circulation fluid. It’s then mixed with a mixture of air and recycled gas, which is compressed by a blower. The gas mixture is pre-heated again by the hot circulated fluid before entering the first reactor. The reactors are fixed-tube reactor filled with iron-molybdenum catalyst. Most of the methanol is converted to formaldehyde in the series of reactors and heat is released from the reactors since the reaction is strongly exothermic. Heat released by the exothermic reaction is removed by vaporization of a high boiling heat-transfer fluid (hot oil) on the outside of the tubes for heat control and hot spot control [5].
The formaldehyde gas is cooled before entering the absorber where it’s either absorbed in water or in diluted urea solution depending on the desired product. The absorbers can also run both parallel and singular.
2.3 Process differences The metal oxide catalyst typically has an effective life of 12-18 months and compared to the silver catalysts it’s more insensitive to contaminants. It therefore requires less frequent catalyst clean-outs but the downtime when replacing the catalyst is longer. The silver catalyst has a lifetime of 3-8 months and is easily poisoned but can be recovered [4].
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The feed of methanol and air to the reactor is a flammable mixture. To avoid the mixture from spontaneous combustion it’s necessary to work outside the explosion limit. To do this, the metal oxide process is working with an excess of air whereas the silver process is working with reduced oxygen content. When working with excess air in the metal oxide process at least 13 mole of air per mole of methanol is necessary [3]. The metal oxide process that works with excess air is working under the lower explosion limit of the mixture. The silver process is working with deficient of air above the upper explosion limit [6].
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3. Method The process will be optimized with respect to heat integration using pinch analysis. The pinch analysis software used in this report is Aspen Plus V8.2 and Aspen Energy Analyzer. The simulation will be the basis for the possible improvements and optimization that could be made.
3.1 Pinch Pinch technology is a method used for minimizing the energy consumption of chemical processes. Matching cold and hot process streams so that the external heating and cooling requirements are minimized. The Pinch technology identifies a “pinch point” which is the minimum driving force allowed in the heat exchanger. This point separates the overall temperature area into two subsystems, one above the pinch point and one below that can be studied separately. External heating must be supplied to the process at temperatures above the pinch point whereas external cooling must be supplied at temperatures below the pinch. There should be no heat transfer across the pinch point [7].
The pinch technology consists of three steps: (1) identification of hot and cold streams, (2) calculation of minimum energy requirements (cooling and heating) and (3) design of heat exchanger network. In the first step properties such as flow rates, phase changes and the temperature range of which must be heated or cooled are identified for the process streams. Hot streams are streams that should be cooled and cold streams are streams that should be heated. The data that is identified is supply temperature, target temperature, heat capacity (Cp), flow rate (m) and enthalpy change (∆𝐻). The enthalpy change is calculated as [7]:
∆𝐻 = 𝑚 ∙ 𝐶! ∙ ∆𝑇 (3.1)
The phases and temperatures of the streams limit the possible heat exchanges between hot and cold streams in the heat exchanger. To make sure that the heat in the streams is fully utilized there are some rules to make sure that the temperature of the streams are lowered by a minimum. This value depends on which phase the stream has. The values used in our analysis can be seen in Table 1. The temperature is lowered to guarantee that the hot and cold streams integrate with each other making sure that the heat is used efficiently.
Table 1 Minimum temperature in heat exchanger.
Gas 50 °C
Liquid 10 °C
Condensing/Evaporating 5 °C
Secondly a “Composite curve” for the process is computed and the pinch point is identified. The composite curve is a temperature-enthalpy diagram, which represents the heat available (hot composite curve) in the process and the heat demands (cold composite curve) in the process. The hot end indicates the minimum heating requirement whereas the cold end indicates the minimum cooling requirement whereas the area where the two curves overlap represent the maximum heat recovery within the process. The minimum temperature difference represents the pinch. The third step consists of designing the heat exchanger networks based on the pinch point location to achieve optimal heat integration [7].
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3.2 Aspen Aspen Plus is software used for chemical process optimization. The features used in this report are the Aspen Plus V8.2 simulator and the Aspen Energy Analyzer.
The Aspen simulation is used to construct, model and optimize the existing process. The plant flow sheet can be exported to the Aspen Energy Analyzer in which the energy integration is evaluated. Aspen Energy Analyzer allows the user to evaluate the energy saving potentials and identify possible changes to the flow sheet to reduce process energy usage. By using this tool ideas to cut down on ineffective energy use and a better energy integration is possible. The Energy Analyzer uses pinch technology to calculate the energy saving possibilities. The composition curve and as well as a suggested heat exchanger network is generated by Aspen Energy Analyzer for the user to identify the best design solutions [8, 9]. From Aspen Plus complete mass- and energy balance are generated.
3.3 Aspen Simulation and modelling When simulating the process based on the flow sheet provided by Haldor Topsøe, several assumptions and simplifications are made when modelling and simulating the process.
3.3.1 Reactors The reactors produce formaldehyde as the following reaction [10]:
𝐶𝐻!𝑂𝐻 +!! 𝑂! ⇌ 𝐶𝐻!𝑂 + 𝐻!𝑂 (3.2)
With the following side reactions [10]:
𝐶𝐻!𝑂 +!!𝑂! ⇌ 𝐶𝑂 + 𝐻!𝑂 (3.3)
2 𝐶𝐻!𝑂𝐻 ⇌ 𝐶𝐻!𝑂𝐶𝐻! + 𝐻!𝑂 (3.4)
𝐶𝐻!𝑂 +!!𝑂! ⇌ 𝐶𝐻𝑂𝑂𝐻 (3.5)
𝐶𝐻! + 𝑂! ⇌ 𝐶𝑂! + 𝐻!𝑂 (3.6)
As can be seen formaldehyde need to be separated which is done in an absorption tower later in the process.
In Aspen the two reactors are simulated to be run in series were the flows in to the first reactor are methanol and air and the product out from the first reactor is going in to the second reactor with another methanol stream. This is to maximize the conversion of the unreacted reactants and will give more formaldehyde as product.
When the reactors were simulated in Aspen, NRTL was chosen as method because the solution into the reactors is non-ideal. No-ideal solutions means that every component in the solution has the same interactions with each other and behaves as it was only one kind of component in the solution. The NRTL as chosen method is also a good approximation for this simulation since the methanol-air mixture in to the reactors are polar and non-electrolytes which according to Figure 2 and Figure 3 as guidelines gives NRTL as method [11].
The simulation of the reaction is run in RStoic-reactors in Aspen. The RPlug would be a better representation of the multi-tube reactors used in the actual plant but since there were unknown kinetic parameters the simulations in these reactors did not give a mass flow that
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well represented the real values. The RStoic reactor is normally used when the reaction kinetics are unknown or unimportant. The product stream is instead calculated using the stoichiometric parameters and the extent of the reaction specified by the user [12].
Figure 2: Guideline for choosing property
Figure 3: Guideline for choosing property method
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3.3.2 Absorption The product out from the reactors contains formaldehyde and this should be separated in an absorption tower. Formaldehyde out from the reactor is in gas phase and when water is added to the absorption tower the formaldehyde could be absorbed to form formalin (formaldehyde in water solution). When a urea-water solution is added in the absorption tower the final product is urea-formaldehyde.
The absorption were simulated as two towers in Aspen were both towers had the water and urea-water solution as inlets streams to in this case only produce urea-formaldehyde with a formaldehyde content of 60 wt%. But since the absorption towers are run as two separated units it is possible to produce both formalin and urea-formaldehyde by adding urea-water solution to only one tower.
The absorption in water is complicated mainly because of two factors, reactions in liquid phase and because it’s exothermic. In aqueous solutions, formaldehyde reacts with water to methylene glycol and higher poly(oxymethylene) glycols according to reaction (3.7) and (3.8) [13].
𝐶𝐻!𝑂 + 𝐻!𝑂 ⇌ 𝐶𝐻! 𝑂𝐻 ! (3.7)
𝐶𝐻! 𝑂𝐻 ! + 𝐻𝑂 𝐶𝐻!𝑂 !!!𝐻 ⇌ 𝐻𝑂 𝐶𝐻!𝑂 !𝐻 + 𝐻!𝑂 (3.8)
These side reactions are implemented in Aspen but all glycols are approximated with methylene glycol because they’re assumed to have similar properties.
Also formaldehyde and free urea can react and form monomethylol urea (equation 3.9), which further reacts to dimethylol urea (equation 3.9) under basic conditions at pH 7-9. This reaction is exothermic [14].
𝐻!𝑁𝐶𝑂𝑁𝐻! + 𝐶𝐻!𝑂 ⇌ 𝐻!𝑁𝐶𝑂𝑁𝐻𝐶𝐻!𝑂𝐻 (3.9)
𝐻!𝑁𝐶𝑂𝑁𝐻𝐶𝐻!𝑂𝐻 + 𝐶𝐻!𝑂 ⇌ 𝐻𝑂𝐶𝐻!𝑁𝐻𝐶𝑂𝑁𝐻𝐶𝐻!𝑂𝐻 (3.10)
The formic acid in the product gas that is passed to the liquid phase must be neutralized by adding a strong base [14]; in this case NaOH is added. The urea-formaldehyde reactions are not considered in the simulation due to time limitations.
Solutions of formaldehyde and water are highly non-ideal. These systems must be modelled using activity coefficient methods, which require binary interaction parameters. When modelling the absorption tower, UNIFAC (Universal Quasi Chemical Functional Group Activity Coefficient Model) is chosen as the thermodynamic model in Aspen. The UNIFAC method predicts the activity coefficients in non-electrolyte liquid mixtures [15]. By using functional group interaction parameters for the molecules that make up the mixture, activity coefficients for the binary and multicomponent mixtures may be predicted. The method is applicable to systems where data is limited, often with good accuracy [16].
The interaction parameter for the urea molecule were calculated in Aspen with UNIFAC as method by splitting up the molecule to its atoms, see structure for urea molecule in Figure 4
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Figure 4: Structure of urea molecule
The urea molecule is drawn in Aspen according to Figure 4 and the interaction between the atoms in the molecule could be calculated. As can be seen the molecule contains a double bound between carbon and oxygen, two single bounds between each nitrogen atom and the carbon, and four hydrogen atoms, were two is bound to each nitrogen atom.
The stages in the tower were assumed to be ideal stages when simulating in Aspen, but in reality the tower contains packed beds. This was made because there was not enough information about the tower.
3.4 Heat transfer coefficient calculations Since the area of the heat exchangers will depend on the energy that is exchanged (QHE) and the temperature difference (ΔT) and also the overall heat transfer coefficient (k) (3.11):
𝑄!" = 𝑘𝐴∆𝑇 (3.11)
Since the overall heat transfer coefficient depends on the heat transfer coefficient (α) these needs to be estimated in order for Aspen Energy Analyzer to be able to determine the correct area of the heat exchangers. The estimation of the heat transfer coefficient can be done for some chemical species in Aspen Energy Analyzer while others need to be estimated or calculated. Since there is hard to determine the exact value of the heat transfer coefficient only approximate values has been calculated and used during the analysis.
All the streams and their name are defined according to Appendix 1.
Certain assumptions were made regarding the heat transfer coefficient (HTC) e.g. for the boiler feed water the value of water from Aspen Energy Analyzer was used. For the streams containing mainly nitrogen and oxygen (Recyclegas ChillAbs, Tailgas, ChillAbs, and Gas to incinerator) the HTC for warm air (100 W/m2/C) was used.
The streams containing a mixture of methanol and air (the feed streams and the stream out of the second reactor) a mean value between the HTC for methanol and air was used.
The HTC-values were the streams did not change phase was calculated through the Dittus-Boelter equation (3.12).
𝑁𝑢! = 0.023𝑅𝑒!!.!𝑃𝑟!.!! (3.12)
𝑅𝑒! ≳ 10000
0.6 ≤ 𝑁𝑢 ≤ 160
Can be rewritten as:
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!∙!!= 0.023 !∙!∙!
!
!.! !∙!!!
!.!! (3.13)
Since this is done for the streams into and out of the absorption towers (Higher Abs 1/2, Lower Abs 1/2, Product Stream 1/2) the estimation made is that the following streams will behave like water since it is a water solution (see Table 2).
Table 2. Values used to calculate the HTC through the Dittus-Boelter equation.1
d (m) µ (kg/(m×s) λ (W/(m×C) ρ (kg/m3) u (m/s)
20×10-3 406×10-6 0.668 977.8 20
The Cp used is the Effective Cp calculated by Aspen Energy Analyzer for each stream (see Appendix 1).
The streams that changed phase from liquid to gas were calculated by the Mostinski equation (3.14). These were the Methanol Feed-stream and the Hot Oil Streams.
𝛼 = 0.1011 ∙ 𝑝!"!.!" ∙!!
!.!1.8𝑝!!.!" + 4𝑝!!.! + 10𝑝!!" (3.14)
!! !
!!∙!!"#= 0.16 !∙!(!!!!!)
!!!
!.!" (3.15)2
!!= !
! !1+ !!
! (3.16)
𝑝! =𝑝𝑝!"
𝑝! = 𝜎 ∙ 𝑔 ∙ 𝜌! − 𝜌!!.!
Below in Table 3 the values that used in the Mostinski equation calculations can be seen.
1 Alveteg M., Handbook, Department of Chemical Engineering Lund University, 2013. 2 Katto Y., Critical Heat Flux, Department of Mechanical Engineering Nihon University, 1994.
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Table 3. Values used to calculate the HTC through the Mostinski equation
Streamname Component ρg (kg/m3)
ρl (kg/m3)
pcr (bar)
p (bar) σ (N/m) Hvap (kJ/kg)
Methanol Feed Methanol 5.64 3 704.0 3 80 3 4.5 22×10-3
1 1100 1
Hot Oil Stream Dowtherm A 4
6.30 828.3 31.34 1 0.0357 799.2
Hot Oil Stream Dowtherm A 4
6.30 828.3 31.34 1 0.0357 799.2
To calculate the HTC for the condensing oil (Cold Oil Stream 1/2) the Nusselt’s Film Theory was used (3.17). The values of the variable used during these calculations may be seen below in Table 4.
𝛼 = 0.943!!∙!∙!!
!∙!!
!!∙!∙!! (3.17)
Table 4. Values used to calculate the HTC through the Nusselt’s Film Theory4
µf (kg/(m×s) λf (W/(m×C) ρ (kg/m3) ΔH (kJ/kg) ΔT (°C) L (m)
0.24×10-3 0.0971 828.3 284.9 15 1.5
The heat-exchange with the condensing oil occurs with different media depending on if the new or old HEN-design is used. It is estimated that the ΔT is not affecting the value of the heat transfer (HTC) constant significantly. Since the HTC calculated are approximate values the difference between different ΔT are neglected and the same ΔT (15 °C) is used for the condensing Downtherm A regardless of what the cooling utility is.
3.5 Calculations for steam turbine In order to produce electricity using a steam turbine the Rankine cycle is used as the idealized model to show the thermodynamic cycle of a heat engine that converts heat into mechanical work, see Figure 5 [24] Water is usually used as the working fluid continuously in the closed loop. The steam is normally superheated before it enters the steam turbine in order to achieve higher cycle efficiency. The conversion of the mechanical work into electricity is achieved with an electrical generator coupled with a shaft to the turbine.
3Methanol.org, “Technical Information”. Retrieved from:http://www.methanol.org/Technical-Information/Resources/Technical-Information/Physical-Properties-of-Pure-Methanol.aspx. Accessed on: 2015-05-05.
4 Dowtherm A, Product Compendium
12
Figure 5: Rankine Cycle
Three different cases are studied and compared where steam at different pressures is produced, 30 bars, 16 bars and 5 bars. The moisture content in the stream that enters turbine should be less than 10% and this is because of higher moisture content gives corrosion and erosion problem in the turbine [17]. Also it is reasonable to assume an isentropic efficiency of 80%. Therefore the discharge pressure and condensation temperature of these three cases can be obtained from the Mollier diagram, see appendix 4 and Table 5. Calculation for the actual enthalpy with 80% isentropic efficiency is using equation (3.18). Notice that the steam in all the cases are superheated to 270 °C base on an oil temperature of 280 °C and a minimum temperature difference of 10 °C in heat exchanger.
𝜂!"#$%&'(!) =!!!!!!!!!!!
(3.18)
Table 5. Steam pressures and condensation temperatures obtained from the Mollier diagram.
Case
Inlet Pressure Discharge Pressure Condensation Temperature
1 30 bar 3.5 bar 142 °C
2 16 bar 1 bar 100 °C
3 5 bar 0.12 bar 50 °C
Following assumptions are used in the simulation of the turbine loop:
Isentropic efficiency 0.8
Mechanic efficiency 0.95
Pump efficiency 0.8
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3.6 Calculations for cost estimation In order to estimate the validity of the results cost estimations were done using a data base method called the Ulrich method to determine the capital cost estimation. This database method considers the general equipment cost, the cost of installation and unforeseen expenses. To see if the investment is good or not the payback time was used in order to see how long it takes for the investment cost to be repaid.
The equations used during the cost estimation can be seen below. The payback time is calculated using equation 3.19 and the capital costs are estimated using equation 3.20.
Payback time:
𝑃𝐵 = −(!" (!!!∙
!!!!)
!" !!!! ) (3.19)
Capital cost estimation, Ulrich method:
𝐶!" = 𝐶! ∙ 𝐹!"! (3.20)
The model factor takes into consideration additional costs needed for the plant and additional cost during extreme conditions e.g. high temperature, high pressure and corrosive environments.
𝐾$,!"#$ = ( (𝐶!" )! )!!!! ∙ 𝑓!""#/!"#$%#&'#!( ∙ 𝑓!"#$%$!&' !"#$%$&$'( (3.21)
𝑓 !""#!"#$%#&'#!(
≈ 1.18 𝑓𝑜𝑟 𝑎 𝐵𝑟𝑜𝑤𝑛𝑓𝑖𝑒𝑙𝑑 𝑃𝑙𝑎𝑛𝑡
𝑓!"#$%$!&' !"#$%$&$'( ≈ 1 𝑓𝑜𝑟 𝑎 𝐵𝑟𝑜𝑤𝑛𝑓𝑖𝑒𝑙𝑑 𝑃𝑙𝑎𝑛𝑡
𝑓 !""#!"#$%#&'#!(
≈ 1.15 𝑓𝑜𝑟 𝑎 𝐺𝑟𝑒𝑒𝑛𝑓𝑖𝑒𝑙𝑑 𝑃𝑙𝑎𝑛𝑡
𝑓!"#$%$!&' !"#$%$&$'( ≈ 1.25 𝑓𝑜𝑟 𝑎 𝐺𝑟𝑒𝑒𝑛𝑓𝑖𝑒𝑙𝑑 𝑃𝑙𝑎𝑛𝑡
This plant is considered a Brownfield plant because the calculations are done on an already designed plant. Greenfield calculations are also made since the process is still sold by Haldor Topsøe and new plants are built.
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4. Results 4.1 Energy optimization The energy optimization is performed through different stages. First evaluation of the existing process, then suggestions of new design is done. Based on the new design two options are suggested and have been studied and compared in regard to energy and economic view.
4.1.1 Evaluation of the existing process All of the heat exchanging streams in the process and the values needed for the calculations were written into Aspen Energy Analyzer and calculated by the program but also by hand. All of the relevant values used can be seen in Appendix 1.
In order to evaluate the existing heat exchanging network (HEN) Aspen Energy Analyzer was used to draw the network. From the values in the table in Appendix 1 from Haldor Topsøe the program calculated the areas of the heat exchangers. The composite curve for the process was also obtained through Aspen Energy Analyzer to analyse the current situation.
In Figure 6, the red line represents the composite curve for the hot stream and the blue line represents the composite curve for the cold streams. From the curve it can be seen that the process requires no external heating because all heating is covered by the internal heat exchange. The process required 3.53 MW external cooling.
From the flow sheet it is calculated that external cooling of 3.14 MW (2.7 Gcal/hr) is added to the process today. The differences in cooling utility that is needed can be due to irregularities in the flow sheet from Haldor Topsøe.
Figure 6: The composite curve for whole process
The HEN as it is today can be seen below in Appendix 2. It has been estimated that cooling water is used as the coolant as far as this is possible. For some of the streams a refrigerant at -25 °C needs to be used. This is not the confirmed condition at the plant but the simulation will
15
be compared with the suggested future design solutions hence a quantitative comparison of savings in energy and heat exchanger area will be presented in the results.
Since there is a significant importance in maintaining a low temperature in the reactors the use of Downtherm A (the oil) as a coolant cannot be used during the development of the new HEN, therefore these streams are neglected during the evaluation. Also the boiler feed water is neglected since it is not part of the formalin production plant but often originates from other nearby production plants.
New composite curves for the process is thereby obtained and shown in Figure 7 and a higher external cooling rate is needed since the neglecting of the boiling feed water. The new cooling rate needed is 4.75 MW according to Aspen Energy Analyzer.
Figure 7: New composite curves for the process
4.1.2 Redesign of the existing heat exchanger network In the given original flow sheet most external cooling utilities are applied on the absorber columns, including the formalin absorbers and chilling absorber. The partial oxidation reaction itself is extremely exothermic and lots of heat is released, therefore no heating utilities are required instead only cooling utilities.
The existing heat-exchanger network in Appendix 3 shows that when streams heat exchange with each other, three of them have crossed the pinch temperatures. This is against the pinch rule where the energy in the system has not been utilized in the best way. According to the composite curve in Figure 7 it’s possible to move the cold composite curve to the left so the streams that can be heated using hot streams apart from the oil. If this high temperature oil is freed it can be used to produce some high quality steam and therefore to design a heat exchanger network has optimal energy utilization.
16
In Figure 8 a simplified picture of the part of the existing HEN that is the focus of the redesign. In Aspen Energy Analyzer the pinch lines has been calculated regarding the process 283 ºC. The heat exchange between the stream from the second reactor and the methanol feed is crossing the pinch line.
Also the heat exchange between the oil stream connected to reactor 1 and the third stream is the gas stream in to the incinerator that crosses the pinch line.
Figure 8: Part of the existing HEN.
The total calculated area of this part of the HEN as it looks today is 105.1 m2. The individual area and temperature for each heat exchanger can be seen in Appendix 3.
The suggested HEN contains High Pressure Steam Generation (HP Steam Generation) as coolant for the condensing oil streams and a heat-exchange loop between the ventilated gas and the gas going in to the incinerator. The need of air blowers for cooling the streams has been eliminated and instead just a little bit of cooling water is needed where HP Steam cannot be produced. The suggested HEN is shown in Figure 9. There is seen that in this HEN as well there is a crossing of the pinch line since the feed out of the second reactor is still used heating the stream into the first reactor.
For the new design the tail gas from the incinerator is cooled to 52 ºC since this gas is only vented out it does not need to have a specific temperature when leaving the process plant.
A simplified picture of the suggested HEN can be seen below in Figure 8. Another solution is having a cooler directly after reactor 2 to make sure the stream that heat exchange with feed to reactor 1 always has same temperature. This is preventing the temperature change from reactor 2 due to the effect of consumption of catalyst.
17
Figure 9:The suggested HEN
Figure 10: A simplified flow sheet of suggested HEN
The total calculated area of this part of the HEN as it looks today is 228.3 m2. The individual area and temperature for each heat exchanger can be seen in Appendix 3.
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4.1.3 Steam and electricity production The aim of the design is to use the energy in the high temperature oil to produce steam and electricity using a steam turbine. There are two options for the produced steam; either just to sell the steam or produce electricity by using a steam turbine in order to cover the energy demand in the plant. The two alternates are studied and will be compared in regard to energy and economic view.
4.1.3.1 Steam The amount of steam available depends on the energy content in the oil, since the heat from the oil will evaporate the water to produce steam. The oil is used to remove heat from the exothermic reaction; this energy is then transferred to different parts of the process. The energy that can be transferred from the oil is the same energy that is used to create the steam. The energy content in the oil loops were taken from the flow sheet, which is provided from Haldor Topsøe, see steam data from Appendix1.
Oil loop 1: 0.79 𝐺𝑐𝑎𝑙/ℎ
Oil loop 2: 1.04 𝐺𝑐𝑎𝑙/ℎ
Total oil energy content is: 1.83 𝐺𝑐𝑎𝑙/ℎ = 2,070 𝑘𝑊
The produced steam flow can vary depending on the pressure of the steam that is desired. Typical useful pressure of steam for exportation is 5 bar, 16 bar, 30 bar and 50 bar. The amount of steam at the different pressures that can be produced is calculated in Aspen by simulating a loop for the steam production.
All the feed water are assumed to have a reasonable inlet conditions that have the same overall pressure through the steam production and the inlet temperature is regulated to keep the feed water below its boiling point. A table over the inlet and outlet conditions and the amount of steam that can be produced is showed below in Table 6.
Table 6: Steam production for 5 bar, 16 bar, 30 bar and 50 bar steam
Inlet
Temperature (°C)
Outlet Temperature (°C)
Steam
Pressure (bar)
Steam Flow (kg/hr)
90 160 5 3,085
150 210 16 3,341
180 240 30 3,500
200 270 50 3,589
4.1.3.2 Steam Turbine A simple flow sheet for the steam-turbine set (Figure 11) is simulated in program Aspen Plus in order to obtain the work that been done in the turbine, see Table 6. In the flow sheet B1 and B2 represent the two heat exchangers where oil boils up the feed water from its condensation temperature to a superheat temperature of 270 °C.
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Figure 11: Steam turbine set
Results for 5 bar, 16 bar and 30 bar are showed in the Table 7. It shows the steam mass flow, work that obtains from turbine, work required for pump and energy released from condensation.
Table 7: Result for steam with 5bar, 16bar and 30 bar.
Pressure (bar) Steam Flow (ton/hr)
Work Turbine
(kW)
Work Pump
(kW)
Condensation Heat ( kW)
5 2.66 368 3.26 1,683
16 2.88 326 1.63 1,729
30 3.12 290 0.463 1,768
As the results table shows, to obtain a maximal work from turbine the 5 bar steam is best to use, the condensation temperature is around 50°C, which may not have many usages.
If higher condensation temperature is desired instead, the steam of 30 bar is better to use. The condensation temperature in 140 °C can be used as heating utility for other nearby plants. If a 16 bar steam is used a condensation temperature of 100 °C is obtained and this can be possible to use both as heating medium for other process plants or as district heating.
According to the results from simulation of the three cases, the work obtained from steam turbine is enough to cover the energy demand in the air compressor and perhaps some of the energy demand for the recirculation pumps in the plant.
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A flow sheet for the turbine loop with 30 bar steam as an example is shown Figure 12.The energy obtained from the turbine is covering the work required for the air compressor.
Figure 12: Flow sheet for turbine loop with 30 bar steam.
4.1.4 Heat and Mass balance The mass balance for those important blocks in the current process obtained from simulations from Aspen plus. The results in Appendix 5 showed total mass and mole flow in and out from every important unit operation.
The heat balance is presented in form of Sankey diagram. In order to compare the current process (Figure 13) and the redesigned one (Figure 14), both Sankey diagram are drawn. There are two mainly differences: firstly the steam production has increased by 4 % and secondly less energy become loses.
21
Figure 13: Sankey diagram for current process: Steam production 10% and Losses 13%.
Figure 14: Sankey diagram for current process: Steam production 15% and Losses 10%.
4.2 Cost analysis To make an economical evaluation of the new design option a comparison of the cost today regarding equipment, installation and possible revenues are estimated. The revenue generated from steam production as well as electricity production will be compared.
The pay-back time is also calculated to evaluate the investment. This plant is a Brown Field Plant, meaning it is an existing plant. All the calculations are based on Brown Field Plants, but there will be some data about Green Field Plant for future reference.
All improvements are dependent on installing a new heat exchanger network, in addition to installing this new network there are some other changes that can be done.
4.2.1 Heat exchanger calculation Shell and Tube heat exchanger is chosen as general heat exchanger type for the HEN design and for calculation of economic fixed tube sheet and U tube is chosen. Reason for the choice is that the shell and tube heat exchanger is the most common type of heat exchanger used in
22
the chemical process industries. It has the lowest cost especially when made of carbon steel. It can also stand high pressure and temperature, 600 °C and 310 bar in the shell and 1,380 bar in the tubes. For detailed design, the fluid inside the tube has much higher pressure than the other used as rules of thumb. [18]
When calculating the cost for the heat exchangers it was assumed that shell and tube heat exchangers with fixed sheet and u-tube as desired subtype was used and that the heat exchangers was made of carbon steel (Cs) as the chosen material. The chemical engineering plant cost index (ICE) was taken for May 2014,which was decided to 574.4. [19]
In the Table 8 and Table 9 below show the cost for each heat exchanger in the current network and suggested network, the cost was calculated using the Ulrich method [20].
Area for heat exchange between air and streams in the process – Aspen cannot calculate for air-heat exchanger (fin-fan) so the area is calculated for a shell and tube. Lesser area may in real life be needed but since this was not possible, we chose to calculate all heat exchangers and their costs as shell and tube.
Table 8 – The costs of the current Heat Exchanger network
Hot stream
Cold stream
Surface [m2]
Total purchased cost [$]
Bare module cost
Cost with Brown Field fees
[$]
Cost with Green Field fees
[$]
Cold Oil Stream 1
Gas to Incinerator
41.3 10,461 33,249 39,234 43,016
From Reactor 2
Air 31.5 9,078 33,249 39,234 42,891
Cold Oil Stream 1
Methanol-Air Feed
9.1 5,123 16,293 19,226 21,018
From Reactor 2
Methanol Feed
3.2 3,764 11,968 14,123 15,439
Cold Oil Stream 1
Air 5.8 4,297 13,664 16,124 17,627
Cold Oil Stream 1
Boiler feed water
14.2 6,197 19,709 23,257 30,002
Total Cost with Brown Field fees [$] 151,198
Total Cost with Green Field fees [$] 169,993
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Table 9 The cost of the new suggested Heat Exchanger network
Difference in cost between new and old heat exchanger network:
156,387− 151,198 = 5,189 USD
Installing the new heat exchanger network is going to cost 5,189 USD more than the old network. But this will enable the heat of the oil to be used for more high quality steam production.
There is a possibility to produce high quality steam and sell it or use it in other nearby process plants. The benefit of this is that no further equipment needs to be installed so there is no big investment cost. But this can only be used if there is a process on the same plant that needs steam since steam cannot be stored and transported a long distance.
As mentioned before the amount of steam available depends on the energy content in the oil, since the heat from the oil will evaporate the water to produce steam.
Total oil energy content is: 1.83 !"#$!= 2070128 𝑊
Assume an 80% efficiency and that the plant operates 330 day a year:
Hot stream
Cold stream
Surface [m2]
Total purchased cost [$]
Bare modul cost
Cost with Brown Field fees
[$]
Cost with Green Field fees
[$]
Cold Oil Stream 2
HP steam generation
63.1 13,218 42,036 49,602 54,226
Cold Oil Stream 1
HP steam generation
31.5 9,078 28,869 34,065 37,241
From Reactor 2
Methanol-Air Feed
9.1 5,123 16,293 19,226 21,018
From Reactor 2
Methanol Feed
3.2 3,764 11,960 14,113 15,428
From Reactor 2
Cooling water
5.8 4,297 13,664 16,124 17,627
Tailgas Gas to incinerator
14.2 6,197 19,709 23,257 30,002
Total Cost with Brown Field fees [$] 156,387
Total Cost with Green Field fees [$] 175,542
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0.8 ∙ 2.07 𝑀𝑊 ∙ 330 ∙ 24 = 13,116 𝑀𝑊ℎ
Since different qualities of steam should have different prices, but the energy content are the same so the steam price is assumed to have approximately 33 USD/MWh in order to simplify the calculations, [21] the price depends on what it replaces, it can be compared to electricity and oil prices.
Total revenue:
33 ∙ 13,116 = 432,828 𝑈𝑆𝐷/year
In this case the investment cost will be the cost for the new heat exchanger network, which is 156,387 USD. The revenue from selling the steam will be 432,828 USD. A short summery of steam cost calculation is shown in Table 10.
Table 10 Summery of steam cost calculation for first year.
New HEN-cost [$] 156,387
Income from selling steam [$/year]
432,828
Total profit first year [$] 276,441
Pay Back time (days) 140 (4.7 month)
4.2.2 Electricity production There is a possibility to install a turbine to produce electricity from the steam produced in the process. The three different pressures are chosen for production of electricity in a non-condensing steam turbine (30 bar, 16 bar and 5 bar) will be studied from an economic view in order to determine which is the most cost efficient.
This can only be done if the new heat exchanger network is also installed. This means that the savings and the costs from the new HEN will be added in the calculations.
An investigation of the investment cost was made to see if the saving each year is big enough to support the total cost. Since an external loop containing a turbine, a condenser and pump are needed to be able to produce the electricity all of these components are considered when calculating the investment cost.
𝐾$,!"#$ = ( (𝐶!" )! )!!!! ∙ 𝑓!""#
!"#$%#&'#!(∙ 𝑓!"#$%$!&' !"#$%$&$'( = 530,428 𝑈𝑆𝐷 [22]
Installation cost for the HEN: 156 387 USD
For calculations on turbine a shaft power of 400kWwas used therefore total investment and installation cost for the turbine set is:
𝐺𝑖 = 530,428+ 156,387 = 686,815𝑈𝑆𝐷
The industrial electricity cost per MWh in Denmark year 2014 was 35.84 USD/MWh [23]. Using Aspen plus to obtain the produced power and the cost estimation methods earlier presented the payback time and investment cost for producing 5, 16 and 30 bars steam were determined and can be seen in Table 11 below.
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Table 11 Comparison of the different steams
Pressure [bar] 5 16 30
Power [kW] 368 326 290
G [$] 686,815 686,815 686,815
a [$] 104,473 92,536 82,288
[%] 9 9 9
PB-time [years] 10.4 12.8 16.1
The most beneficial steam to use would be the one at 5 bar if only wanted the most amount of electricity produced. However the payback time is very long for all of the options indicating that this is not a beneficial investment.
4.2.3 Sensitivity analysis The profitability will be investigated by changing the electricity and steam price to see how sensitive the investment is to changes.
4.2.3.1 Electricity production The industrial electricity price will be changed; it will increase and decrease with about 25%.
Figure 15: Total Profit as a function of electricity price
26
The electricity price will vary between 25 to 45 USD/MWh. As can be seen the profit depends on the electricity price. The profit is the difference between the total investment cost, K$ 2015, and the total income from the investment. Since the steam is the main source of income, the steam price will affect the profit, the different steam generates different profit 5 bar steam being the most profitable.
To investigate which parameter that affects the payback time the most, some parameters will be changed. The following graphs will present the results.
The industrial electricity price will be changed; it will increase and decrease with 25%. The industrial electricity cost per MWh in Denmark year 2014 was 35.84 USD/MWh. In the following graph it will vary from 25 to 45 USD/MWh. The graph also shows the difference in profit between the different steam types.
Figure 16: PB-time as a function of the electricity price
The graph shows that the electricity price has some impact on the PB-time, the higher the price is, the better the investment will be and the PB-time will be shorter. But for the steam at 30 bar this did not apply. The PB time does not change linearly with the electricity price.
When calculating the PB-time an internal rate of return at 9% was assumed, this will also be varied to determine how much it affects the PB-time.
𝑖! = 9%
The internal rate of return varies from 1 to 11%.
27
Figure 17: PB-time as a function of the internal rate of return
The graph shows that the internal interest is an important parameter, especially for the steams at 30 and 16 bar. Yet again the 5 bar steam is the most profitable and least sensitive.
28
4.2.3.2 Steam production and selling
The steam price has been varied from 20 USD/MWh to 40 USD/MWh to see how it affects the payback time. The higher the steam price will be, the lower the payback time will be. There has not been any consideration for the different prices of the different types of steams.
Figure 18: Pay-back time as a function of steam prices
29
Figure 19:The profit as a function of the steam price
The correlation between the profit and the steam price follows the relationship between the pay-back time and the Steam price. The higher the steam price is, the higher the profit from selling it.
30
5. Discussion During the analysis a lot of assumptions were made, mostly in order to proceed with the analysis since there were a lot of unknown parameters. The assumptions made that will affect the results the most is that the area of the air coolers used today could not be calculated and were instead estimated as shell and tube heat exchangers. This will make the calculations based on the area of the heat exchangers unreliable since air coolers does not work or look like shell and tube-exchangers. This may affect the difference in the cost between the old and the new heat exchanger network.
Since most of the analysis made in Aspen Energy Analyzer was made from the flow sheet provided by Haldor Topsøe, which may not be completely in mass balance because this types of measurements often are made at different times. This will give a small inaccuracy to our results because the given flows may not be precisely represented by the given cooling-demands in the flow sheet. This will affect the areas of the heat exchangers and therefore also the costs and the amount of energy that can be produced. Although the results from our analysis will give a good representation if the investments are beneficial or not since we assume that these errors are rather small. And also this is just one size of the plant, there are many different sizes of this formaldehyde plant process which may affect the investment costs as well as the profit.
Another thing that could have affect our results are the assumptions for the absorption, where we assumed that there were 20 stages and the different inlet flows of water and urea in to the tower.
Due to the assumptions that were made we obtained a lower content of formaldehyde in the final product of 46 wt-% instead of 60 wt-% that was given in the flow sheet. We believe that the main reason for this deviation is the flow sheet the calculations were based on. As mentioned earlier; to get an accurate mass balance in a flow sheet, the flows and other data must be sampled at different times to get complete information of the process. This gives an accurate mass balance thus an accurate simulation.
Another thing that will affect the cost estimations is the fact that the steam price used was only for one type of steam. The price between low quality and high quality steam tends to vary a lot, and also the steam price in different part of the world varies a lot. This will affect how beneficial the steam production will be and also if the investment of extending the heat exchanger area is a valid one to make. It may be beneficial in some countries but since Haldor Topsøe is selling this plant throughout the world it may not be beneficial everywhere.
There is only one suggested HEN-design resulting from this report and this will increase the heat-exchanger area needed but it will take care of the excess heat in this process which is not done today. There has been two suggestions to how the heat may be used – either for steam-production or to produce electricity. There are advantages and disadvantages with both of these suggestions.
As shown in the results the income from producing electricity is very low and it also requires that a major investment is made. Therefore it has a long payback time, about 10 to 16 years, which will not be seen as a good investment to make by most companies. But electricity is something that is needed everywhere and the need for it keeps growing in the world. It can
31
also be transported from the plant if it is not needed in the nearby area and may therefore be sold in a wider market.
The steam production is much more beneficial if we just look at the economical view used in this report. It does not require an extra investment besides the new heat exchanger network and the steam produced can be high quality steam resulting in a large income. But the truth is that the steam price may vary a lot in different parts of the world resulting in a longer payback time in some countries. Also the steam needs to be sold in nearby plants or used in the own production since it cannot be stored or transported long distances. This may be a problem if the company owning this facility does not have other nearby production plants or if it is located in an area where the steam price is very low due to over production.
It has also been pointed out to us that the stream out of the second reactor will increase in temperature as the catalyst ages. This means that in order to have an heat-exchange between this stream and the feed-streams to the first reactor the stream either needs to be split and one part cooled separately or a regulated heat exchanger needs to be installed before this heat exchange can be made. This will increase the heat exchanger area needed and therefore also the investment cost for the suggested HEN-design resulting in a slightly smaller profit.
Further investigations can be made in order to obtain a better estimation of the profit if we invest in a new heat exchanging network for example we have only looked at shell and tube heat-exchangers. There could be more beneficial to use other exchangers at least for cooling some of the streams. We have not been looking into how other exchangers may affect the investment cost and heat exchanger area. Also the cooling media used has been mainly cooling water. Other cooling media may require a smaller area of the heat exchanger. And also different plants may use different cooling utilities making it easier for some to use cooling water and for some other cooling mediums may be easier to use.
If the simulation in Aspen Plus were further developed to better represent the flows in the Haldor Topsøe flow sheet this simulation could easily be altered to represent different sizes of plants. This could lead to better representations of how well this HEN would work in different sized plants.
Also in this report only one heat exchanging network (HEN) was suggested since it was considered to be the most reasonable one for the plant. Aspen Energy Analyzer can generate more suggestions that could be investigated if more time is provided. And then a wider range of suggestions may be presented taking both different heat exchangers, cooling utilities and steam/electricity prices into consideration.
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6. Conclusion The suggested new heat exchanger network has better energy utilization, where energy in tail gas has been utilized and air cooled-off energy from the old network has been used to produce steam in the new one. The excess heat would be enough to produce about 13,000MWh steam per year, which is 4% more than the old system.
This heat exchanger network was studied with two options for the produced steam, either selling the steam or using it in to produce electricity using a steam turbine. Both ideas are affected by the electricity and steam prices, but mostly linear affection. The calculations showed that installing a steam turbine would not be profitable at all due to its expansive installation cost; the high investment cost is hard to re-earn giving a payback time ranging between 10-16 years depending on which pressure of steam that is used. Selling steam would generate an income of approximately 270,000 USD the first year and 400,000 USD remaining years, which is more profitable and the payback time is around 5 month.
33
7. References 1. Haldor Topsøe,”History”, http://www.Topsøe.com, retrieved from
http://www.Topsøe.com/about/history. Accessed on 2015-02-16
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34
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I
Appendix 1 Stream data for the current formaldehyde process plant.
Stream name Tin
(ᵒC) Tout
(ᵒC) MCp (kj/(s×ᵒC)
Enthalpy (MW)
HTC (W/(m2×ᵒC))
Flowrate (kg/h)
Effective Cp (kJ/(kg×ᵒC))
ΔT (ᵒC)
Boiler Feed Water
90 158 17.79 1.210 2555.1 1884 34.001 10
Methanol Feed
80 108 10.80 0.3024 16000 857 45.368 10
Methanol-Air Feed
108 200 3.67 0.3373 550 12370 1.067 50
From Reactor 2
285 130 6.828 1.058 550 14530 1.692 50
Recirk ChillAbs
30.9 20 50.15 0.5466 100 55530 3.251 50
Recyclegas ChillAbs
64.9 47.3 44.93 0.7908 100 13390 12.082 50
Gas to incinerator
20 250 1.669 0.3838 100 4553 1.319 50
Cold Oil Stream 1
280 279 860.6 0.8606 710 - - 5
Cold Oil Stream 2
280 279 1210 1.210 710 - - 5
Hot Oil Stream 1
279 280 883.9 0.8839 22400 - - 5
Hot Oil Stream 2
279 280 1244 1.244 22400 - - 5
Reactor 1 286 285 883.9 0.8839 550 12760 249.337 5
Reactor 2 286 285 1244 1.244 550 14530 308.319 5
Higher Abs 1 71.2 65 48.77 0.3024 9020 59920 2.930 10
Higher Abs 2 71.2 65 48.77 0.3024 9020 59920 2.930 10
Lower Abs 1 78 71 43.20 0.3024 8738.9 58410 2.662 10
II
Lower Abs 2 78 71 43.20 0.3024 8738.9 58410 2.662 10
Product Stream 1
78 37 1.135 0.04652 8797 1504 2.716 10
Product Stream 1
78 37 1.135 0.04652 8797 1504 2.716 10
III
Appendix 2
The HEN used today in the formalin production plant by Haldor Topsøe.
Hot stream Hot T in (ᵒC)
Hot T out (ᵒC)
Cold stream
Cold T in (ᵒC)
Cold T out (ᵒC)
Enthalpy (MW)
Area (m2)
Cold Oil Stream 1
279.2 279.0 Air 34.2 35.0 0.1395 5.94
From Reactor 2
240.7 130.0 Air 30 34.2 0.7559 55.93
Product Stream 2
78 37 Cooling water
22.3 22.4 0.0465 0.576
From Reactor 2
286 285 Hot Oil Stream 2
279 280 1.244 388.15
From Reactor 2
285.0 240.7 Methanol Feed
80 108 0.3024 3.396
Lower Abs 1
78 71 Cooling Water
20 20.7 0.3024 2.132
Higher Abs 1
71.2 65 Cooling Water
24.3 25 0.3024 2.631
Reactor 1 286 285 Hot Oil Stream 1
279 280 0.8839 275.71
Cold Oil Stream 1
280 279.1 Gas to incinerator
20 250 0.3838 41.329
Recirk ChillAbs
30.9 20 Refrigerant 1
-25 -24 0.5466 118.32
Lower Abs 2
78 71 Cooling Water
20.72 21.44 0.3024 2.1604
Higher Abs 2
71.2 65 Cooling Water
21.56 22.28 0.3024 2.4756
Cold Oil Stream 2
280 279 Boiler Feed Water
90 158 1.210 14.228
Cold Oil Stream 1
280 279.2 Methanol-Air Feed
108 200 0.3373 9.0865
Product 78 37 Cooling 21.44 21.56 0.04652 0.55817
IV
Stream 1 Water
Recyclegas ChillAbs
64.9 47.3 Cooling Water
22.39 24.28 0.7908 254.08
Total area: 1176.7
Appendix 3
Table 12: The current HEN with the area for each heat-exchanger.
Hot stream Hot T in (ᵒC)
Hot T out (ᵒC)
Cold stream Cold T in (ᵒC)
Cold T out (ᵒC)
Area (m2)
Cold Oil Stream 1
280 279 Boiler Feed Water
90 158 14.2
Cold Oil Stream 1
280 279.1 Gas to incinerator
20 250 41.3
Cold Oil Stream 1
280 279.2 Methanol-Air Feed
108 200 9.1
V
Cold Oil Stream 1
279.2 270 Air 30 31.3 5.8
From Reactor 2
285 217.8 Methanol Feed 30 108 3.2
From Reactor 2
217.8 130 Air 31.3 35 31.5
Total area: 105.1
Table 13: The suggested HEN with area for each heat-exchanger.
Hot stream Hot T in (ᵒC)
Hot T out (ᵒC)
Cold stream Cold T in (ᵒC)
Cold T out (ᵒC)
Area (m2)
Cold Oil Stream 2
280 279 HP Steam Generation
249 249.6 63.1
Cold Oil Stream 1
280 279 HP Steam Generation
249.6 250 45.6
From Reactor 2
285 210.1 Methanol-Air Feed
108 200 15.5
From Reactor 2
210.1 142.9 Methanol Feed 30 108 5.8
From Reactor 2
142.9 130 Cooling Water 20 25 1.1
Tailgas 470 51.6 Gas to incinerator
20 250 97.2
Total area: 228.3
VI
Appendix 4 Mollier diagram with entropy, enthalpy, temperature and pressure. Diagram shows how much energy obtains when the input pressure is 5, 16 and 30 bar.
VII
Appendix 5 Table 14 Mass balance for the simulated process
Total
Mass Flow(kg/hr)
Total
Mole Flow(kmol/hr)
Reactor 1 IN 13 226,83 463,5659
UT 13 226,83 475,6084
Reactor 1 IN 14 511,83 515,7119
UT 14 511,83 535,9202
Air Compresor IN 12 369,83 436,8199
UT 12 369,83 436,8199
Incinerator IN 5701,564 201,4756
UT 5701,642 202,4867
Chilling Absorber
IN 13 117,99 483,8184
UT 11 403,13 402,9512
Excess water 1714,861 80,86723
Absorber1 Feed in 7255,916 267,9601
Product 1366,188 48,76785
Gas out 6572,228 242,4654
Water 104 5,772877
Urea 578,5 17,50035
Absorber2 Feed in 7255,916 267,9601
Product 1392,656 49,88035
Gas out 6545,761 241,353
Water 104 5,772877
Urea 578,5 17,50035
