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<ul><li><p>1 </p><p>DESIGN OF METHANOL PLANT </p><p>EURECHA Student Contest Problem Competition 2013 </p><p>Chan Wei Nian, </p><p>Fang You, </p><p> Department of Chemical &amp; Biomolecular Engineering, National University of Singapore, Engineering </p><p>Drive 4, Singapore 117576, Republic of Singapore </p><p>EXECUTIVE SUMMARY </p><p>This report presents a techno-economic analysis of setting up a methanol plant with a capacity of </p><p>50,000 MT/yr from natural gas feedstock of 31,000 MT/yr and 90% methane. This is performed in the </p><p>European context with due considerations for environmental sustainability. The plant is designed </p><p>assuming 10-year plant life and 8000 hours of operation per annum. </p><p>The methanol process was developed according to open literature and industry standard, as described in </p><p>Section 2; it consists of 5 units/sections: steam methane reformer, compression trains, methanol </p><p>synthesis reactor, hydrogen separation and methanol distillation. These were grouped into two main </p><p>segments of steam methane reforming and methanol synthesis, and were independently optimized in </p><p>Sections 3.1 and 3.2 using Aspen HYSYS v7.2 by varying important process parameters within </p><p>reported ranges. This is followed by a plant-wide heat integration study with Aspen Energy Analyzer </p><p>v7.2, described in Sections 3.3 and 4.3. Operating conditions favouring better overall process profit of </p><p>the order of USD$106 are reported in Sections 4.1 and 4.2 for the ideal case ignoring operating </p><p>constraints, and the base case for economic analysis was chosen based on practical industry constraints; </p><p>the reforming section utilizes an excess steam to carbon ratio of 2.7, inlet pressure of 2000 kPa and </p><p>furnace outlet temperature of 900C and while the methanol synthesis loop employs a reactor </p><p>temperature of 255C, compressor discharge pressure of 8010 kPa and a stoichiometric ratio of 9.035, </p><p>as summarized in Section 4.4. </p><p>The fixed capital investment, cost of manufacturing and revenue for the base case is 31,562,300, </p><p>29,520,600 and 43,952,670 respectively. The payback period is 3.2 years and the net present value </p><p>(NPV) of the investment is 32,840,270. The breakeven price for methanol is 0.3285/kg and the </p><p>discounted cash flow rate of return is 26.02%. The former is most sensitive to natural gas, hydrogen </p><p>and methanol prices, as shown in Section 4.5. In Section 4.6, Monte Carlo simulations performed </p><p>according to expected price fluctuations of natural gas and methanol obtained in the past 10 years </p><p>reported a 7.5% probability of making an overall investment loss. </p><p>The feasibility of substituting biogas for natural gas was analyzed for the given feedstocks in Section </p><p>4.7. Several options were explored, and the highest savings in net present value of 1,299,770 as well </p><p>as reduction of 2074 MT/yr of CO2 emissions were obtained for the case of utilizing the non-liquid </p><p>feeds. Options to capture carbon emissions from the plant as well as other carbon sources were </p><p>explored, and the most suitable technology for the plant was found to be absorption using mono-</p><p>ethanol amines, as elaborated in Section 4.8. </p></li><li><p>2 </p><p>1. INTRODUCTION </p><p>Methanol demand in the global market is burgeoning with a 23% increase from 2010 to 2012 to 61 </p><p>million tons and an expected increase to 137 million tons in 2022 [1]. However, its raw material: </p><p>natural gas faces both supply limitation and volatility; as a non-renewable resource, natural gas is </p><p>expected to deplete as early as 37 years depending on the level of conservativeness [2] while gas </p><p>infrastructure in Europe is fragmented and inconsistent [3]. At the same time, environmental concerns </p><p>for carbon emissions necessitate the integration of carbon capturing technologies. These usher a new </p><p>age of sustainable design in this age-old process, and it will be reviewed in this report. A base case for </p><p>the methanol synthesis was developed and optimized, followed by case studies on biogas substitution </p><p>and alternative carbon capture opportunities. </p><p>2. PROCESS DESCRIPTION </p><p>For the base case, the methanol synthesis is generally split into steam reforming, methanol synthesis, </p><p>hydrogen separation and methanol purification, which are described in this section. </p><p>Steam reforming: This is an overall endothermic reaction involving catalytic conversion of methane </p><p>and steam into syngas at high temperatures above 800oC [4], medium pressures of 15-30 bar [5] and </p><p>steam to carbon (S/C) ratio of 2 to 4 using nickel catalysts. Steam reforming is limited by equilibrium </p><p>[6], and higher conversion is achieved by increasing temperature, lowering pressure and higher S/C </p><p>ratio. It usually occurs in nickel catalyst-packed tubes located in the radiant section of a furnace. </p><p>Steam Reforming: (H25C = 206kJ/mol) ----------------- (1) </p><p>Water Gas Shift: (H25C = -41kJ/mol) ----------------- (2) </p><p>Methanol Synthesis: Syngas can be catalytically converted to methanol via an overall exothermic </p><p>reaction at medium temperatures of 210-270C and high pressures of 50-100 bar [7], over copper-</p><p>alumina catalysts. The presence of the water-gas shift reaction necessitates the use of a modified </p><p>stoichiometric ratio defined below, and it is typically in the range of 9-10 [4]. </p><p>Methanol Synthesis: (H25C = -91kJ/mol) --------------- (3) </p><p>Modified Stoichiometric Ratio: </p><p> -------------------------------------------------------- (4) </p><p>Methanol synthesis reactors are designed to remove reaction heat via cooling service fluids such as </p><p>unheated reactants or boiler feed water as well as feed quenching [8]. Isothermal cooling in shell and </p><p>reactor tube setups was modeled under conditions known to approach equilibrium [4], and it can </p><p>achieve higher conversions with lower temperatures, higher pressures and stoichiometric ratios. </p><p>Methanol Purification: Methanol synthesis products containing methanol and syngas are flashed to </p><p>separate unconverted light ends and crude methanol. Crude methanol is distilled to separate remaining </p><p>light-ends, methanol at 98%wt and water, in an atmospheric column with a partial condenser. </p><p>Hydrogen Separation: The flashed light-ends contain excess hydrogen; to avoid recycle </p><p>accumulation, these must be purged for furnace fueling or purified for hydrogen credit via pressure </p><p>swing adsorption, cryogenic distillation or membrane separation [9] with the latter being cheaper at the </p><p>low recoveries and purities required [10]. For this, polyimide membranes are fabricated as hollow fiber </p><p>tubes in a shell, and separation is achieved by partial pressure differences. </p></li><li><p>3 </p><p>3. METHODOLOGY </p><p>The process was simulated using Aspen HYSYS v7.2, and individual units were designed to sufficient </p><p>detail for costing based on the module factor approach [11]. Expected revenue from methanol </p><p>(0.413/kg [12]) and hydrogen (8.04/kg [13]) sales was factored in for a plant-wide optimization. The </p><p>methodology for design and costing for each unit for the optimization is elaborated below. </p><p>3.1 Steam Reforming Section Optimization </p><p>This section focuses on the reformer furnace with the reaction modeled to reach equilibrium and the </p><p>costing based on energy used. This is independent of downstream synthesis and is optimized </p><p>separately. The variables affecting this section are the pressure, furnace temperatures (i.e., pre-heat and </p><p>outlet temperatures) and S/C ratio; these affect costs of individual sections/units, and their expected </p><p>effect on profit is shown in Table 1. </p><p>Table 1: Expected effect of increase in design variables on profit due to costs of individual sections/units </p><p>Design Variable Energy Recovery </p><p>from Natural Gas </p><p>Convection </p><p>Section </p><p>Radiant </p><p>Section Compression Conversion </p><p>Pressure Uncertain Uncertain Pre-heat Temperature NA Uncertain Outlet Temperature NA NA Uncertain S/C Ratio NA </p><p>Energy Recovery: Natural gas is expected to be delivered at 75 bar within European pipelines [3] and </p><p>must be expanded via a valve or turbine for energy recovery to the desired pressure; the latter can be </p><p>achieved with radial gas turbines at 75-88% efficiency [14] together with a minimum motor generator </p><p>efficiency of 95% mandated by EU Directive 640/2009 [15]. The turbine was priced based on energy </p><p>recovery rate and capital costs. </p><p>Steam Reformer Furnace: Natural gas is mixed with steam and preheated in the convection section </p><p>before reaction in the radiant section. The duties are summed, and used for furnace capital and energy </p><p>costs. Catalyst life is assumed to be within industrial norm, and is priced based on USD$0.55 per kilo </p><p>mole of natural gas processed [16]. A 95% approach to equilibrium was found to best match industrial </p><p>data from [17], and this was used in the reactor simulations </p><p>Water Let-Down Vessel: Excess steam is condensed and removed from the reformer products, in a </p><p>knock-out vessel to ensure dryness for the compression train. The design is based on the water </p><p>throughput to allow for 7.5 min of residence time, and the vessel volume is used for capital costing. </p><p>The condensate is pumped up to process pressure and reheated as recycle steam to the reformer. </p><p>Make-up Compressor: A separate compression train for syngas products (i.e., fresh feed) is designed </p><p>because it has less variability than the recycle compressor and so requires less capacity allowance. For </p><p>the discharge pressure of 50-100 bar and required flow of about 0.410 m3/s, reciprocating compressors </p><p>are suitable and are used for costing at 75% efficiency. To avoid adiabatic temperature rise beyond the </p><p>maximum temperature of 480 K [14], two-stage compression is used with inter-stage cooling. This </p><p>cooler is designed using heat transfer area obtained using typical overall heat transfer coefficients and </p><p>log mean temperature difference for cooling water service [18]. </p></li><li><p>4 </p><p>3.2 Methanol Synthesis Section Optimization </p><p>Due to low equilibrium conversion to </p><p>methanol, recycle of syngas is necessary. The </p><p>effect of reactor temperature, pressure and </p><p>stoichiometric ratio of reactants on </p><p>profitability of this section is shown in Figure </p><p>1. For modeling and optimization purposes, </p><p>common reaction conditions of 210-270C and </p><p>5-10 MPa were used. Costs involved for </p><p>cooling reactant products in the crude </p><p>methanol flash, pre-heating for distillation and </p><p>column duties are calculated based on a </p><p>preliminary setup, and optimized later using </p><p>pinch analysis together with the steam </p><p>reforming heat exchangers. </p><p>Figure 1: Schematic for inter-linked effects of varying </p><p>parameters independent variables (black), parameters affecting cost (shaded pink), calculated parameters (blue) </p><p>Methanol Synthesis Reactor: The recycle and fresh syngas streams are heated to the reaction </p><p>temperature with reaction heat and an optional pre-heating or cooling; this is possible in shell and tube </p><p>reactors [19]. The reactor was sized based on heat transfer area [20]. Catalyst cost, based on its </p><p>replacement norms, is USD$0.0513 per kilo mole methanol, for isothermal reactors [4]. </p><p>Methanol Purification: The crude methanol flash vessel was sized similar to the water let-down </p><p>vessel, and the distillation column was optimized based on typical feed of 78-80%mol methanol at </p><p>70oC. Trade-off between reflux ratio and number of stages was analysed by costing the column, </p><p>condenser, reboiler, reflux drum and reflux pump, and the optimum number stages was found to be 27 </p><p>for reflux ratio from 0.4 to 0.6 (Figure 2). This design is deemed optimal and used for all runs in the </p><p>optimization study. </p><p> Figure 2: Local optimization for distillation </p><p>column </p><p> Figure 3: Hollow Fiber Membrane Separation Module </p><p>Source: [21] </p><p>Hollow Fiber Membrane Module: Hydrogen separation is as shown in Figure 3 to recover hydrogen </p><p>from the purge stream. Feed to permeate pressure ratio of 6 was used to minimize recompression of </p><p>hydrogen [9], and this sets the limit for hydrogen recovery to 95%. The annualized capital cost increase </p><p>is calculated to be in the range of USD$2,442.9 for a 0.01 fraction increase in recovery, whereas the </p><p>gain in revenue was USD$365,600 for the same increase. As such, hydrogen recovery was maximized </p><p>at 95% for each optimization run. With the hydrogen flux constant, only the membrane cost will vary </p><p>1.51 </p><p>1.52 </p><p>1.53 </p><p>1.54 </p><p>1.55 </p><p>1.56 </p><p>20 25 30 35 40 45 50 </p><p>An</p><p>nu</p><p>aliz</p><p>ed</p><p> Co</p><p>sts </p><p>(USD</p><p>$m</p><p>il) </p><p>Number of Stages </p></li><li><p>5 </p><p>according to required membrane area (USD$21/m2 [22]) obtained from changes in log mean pressure </p><p>difference as a result of changing reactor pressure levels for each run. </p><p>Recycle Compressor: The recycle compressor was designed for approximately 0.495 m3/s of flow with </p><p>reciprocating compressors as the suitable choice [14]. Due to a smaller pressure difference, single stage </p><p>compression is sufficient, and the compressor was priced using duty and 75% efficiency. </p><p>3.3 Heat Integration </p><p>In order to improve the energy efficiency of the methanol plant, heat integration is performed to </p><p>recover process heat, using Aspen Energy Analyzer (AEA) v7.2 coupled with the Aspen HYSYS v7.2 </p><p>simulation of the methanol process. The Utility Composite Curve from pinch analysis is shown in </p><p>Figure 4 with minimum approach temperature (dTmin) contribution for various streams listed in Table </p><p>2. From Figure 4, the shifted process pinch temperature is 135.3C and the overall heating and cooling </p><p>targets are 15.9 MW and 13.3 MW respectively. The area target returned by AEA is 2345 m2 for 1 shell </p><p>pass and 2-tube pass heat exchangers. </p><p>Due to the lower price of natural gas compared to steam, AEA program recommends generating high </p><p>amounts of steam from the fired heater as seen in Figure 4. However, the credit from exporting steam </p><p>may be diminished if there is lack of demand for steam. Hence, the plant will not generate excess steam </p><p>from natural gas. </p><p> Figure 4: Utility Composite Curve for Methanol Plant Process </p><p>Table 2: dTmin Contribution for Each of the Stream Types </p><p>Stream Type dTmin contribution (K) </p><p>Condensing/vaporizing 2.5 </p><p>Liquid 5.0 </p><p>Gas 7.5 </p><p>Subsequently, a Heat Exchanger Network (HEN) is designed with the following operating constraints: </p><p>(1) HP steam for SMR is to be generated in the furnace to ensure a steady supply of feed for the </p><p>process; (2) the reforming heat of reaction must be supplied by a furnace to maintain optimal reaction </p><p>temperature as process heat exchange potentially introduces fluctuations; and (3) crude methanol </p><p>distillation column condenser and reboiler are to be serviced by utilities to ensure controllability of </p><p>column operations. </p></li><li><p>6 </p><p>4. RESULTS AND DISCUSSION </p><p>Simulations were run for variations in operating parameters beyond the practical constraints applied in </p><p>industry; this was to explore the potential cost savings from surpassing operating constraints. Data was </p><p>extracted and costed according to Section 3 with detailed calculations in Section A4.2. The optimal </p><p>process was then chosen after applying the appropriate constraints and the heat integration study was </p><p>condu...</p></li></ul>