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Conceptual design of coal to synthetic natural gas (SNG) process based on BGL gasifier: modeling and techno-economic analysis Yang Liu, Yu Qian, Huairong Zhou, Honghua Xiao, and Siyu Yang Energy Fuels, Just Accepted Manuscript • DOI: 10.1021/acs.energyfuels.6b02166 • Publication Date (Web): 25 Nov 2016 Downloaded from http://pubs.acs.org on December 5, 2016
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Energy & Fuels
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Conceptual design of coal to synthetic natural gas
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(SNG) process based on BGL gasifier: modeling and
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techno-economic analysis
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Yang Liu, Yu Qian, Huairong Zhou, Honghua Xiao, Siyu Yang*
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School of Chemical Engineering, South China University of Technology, Guangzhou,
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Guangdong 510640, P.R. China
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+ For publication in Energy & Fuels
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*Corresponding author:
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Professor Siyu Yang Ph.D.
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School of Chemical Engineering
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South China University of Technology
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Guangzhou, 510640, PR China.
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Tel.: +86-20-87112056, +86-18588887467
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Email:
[email protected] 24
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Abstract
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Developing coal to synthetic natural gas (SNG) projects in China could alleviate
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the conflict between supply and demand of natural gas. Coal to SNG process based on
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Lurgi gasifier (Lurgi process) has been commercially implemented. However, low
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energy efficiency and high production cost restricted expansion of Lurgi processes.
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BGL gasifier developed based on Lurgi gasifier is considered to be more appropriate
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for coal to SNG projects.
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In this work, a novel BGL gasifier based coal to SNG process (BGL process) is
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proposed and simulated based on rigorous kinetics models. Based on simulation,
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techno-economic analysis and comparison is conducted over conventional Lurgi
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process. The results show that objective energy efficiency of the BGL process is 8%
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larger than that of the Lurgi process, while increase of general energy efficiency is not
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obvious. Considering the income of product, ROI (return on investment) of the BGL
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process is 3% higher than that of the Lurgi process. However, the Lurgi process can
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produce more byproducts. If considering values of these byproducts, the BGL process
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will not has economic competitive.
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Keywords: coal to SNG, BGL gasifier, modeling, techno-economic analysis
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1. INTRODUCTION
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Natural gas (NG) has long been regarded as one of the most clean energy sources.
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In China, with the requirement of reducing of greenhouse gas emission and the
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awareness of environmental protection, the demand for natural gas is growing rapidly,
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reaching more than 191 billion m3 in 2015.1 According to the IEA forecast, the
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demand of natural gas will reach 360 billion m3 in 2020, which is roughly twice that
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in 2015.2 However, the current conventional natural gas production rate cannot meet
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the continuous growth of demand. The increasing demand has led researchers to
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consider alternate methods of natural gas generation. Converting coal to natural gas,
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namely synthetic natural gas (SNG), is considered as an effective way to satisfy the
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demand for natural gas, while utilizing China’s abundant coal resources.3 Currently,
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there is a commercial-scale demonstrative coal to SNG projects under production in
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the United States, and three in China. Moreover, 54 coal to SNG projects are under
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plan stage with a total capacity of 225 billion m3 annually. Coal to SNG project is
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playing an increasing important role in the aspect of China’s natural gas supply.4
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The key device of coal to SNG process is gasifier. There are three types of coal
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gasifiers: moving bed, entrained flow, and fluidized gasifiers. As a kind of moving
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bed gasifier, Lurgi gasifier is used widely in coal to SNG project, such as Great Plain
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Synfuels plant, Xinjiang Qinghua project, and Datang Keqi project.5 However, these
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projects suffer from low energy efficiency and high production cost.6 To solve these
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problems,
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hydro-methanation (bluegas) process that carries out coal gasification and
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methanation in one apparatus. It is reported that the bluegas technology has higher
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energy efficiency. However, there is no precedent for its commercial-scale
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applications until now. Many engineers and researchers also recommend to use other
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commercial gasifiers to replace Lurgi gasifier. Yu and Chien established a coal to
the
Great
Point
Energy
company
has
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developing
the
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SNG process based on water slurry gasifier and made techno-economic analysis for
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the process.7 Nevertheless, because of high energy consumption and huge capital
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investment of gasifier, both total capital investment and operation cost of this process
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are much higher than Lurgi process.
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BGL gasifier is an extension of Lurgi gasifier with the ash discharge under
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slagging conditions. It has the advantages of both much lower steam and somewhat
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lower oxygen consumption.8 It is considered to be more appropriate for coal to SNG
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project. Currently, BGL gasifier has been successfully applied to 150 thousand tons of
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coal to dimethyl ether plant by Yunnan Coal Chemical Industrial Group and 18
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thousand tons of ammonia plant by Datang Group. These projects provide a basis for
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successful utilization of BGL gasifier in SNG process.
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The objective of this paper is to develop a novel BGL gasifier-based coal to SNG
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process. Specifically, the novel process includes air separation unit (ASU), coal
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gasification (CG) unit, water gas shift (WGS) unit, acid gas removal (AGR) unit, and
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methanation unit. All of these units are modeled in detail models using Aspen plus
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simulator respectively. Among these units, BGL gasifier-based coal gasification unit,
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which has not been reported in the open literature, is systematically modeled in this
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paper. On the basis of detailed process modeling and simulation, a techno-economic
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analysis is conducted to examine the advantages and disadvantages of the novel
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process compared with conventional process.
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2. PROCESS MODELING AND SIMULATION
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The flow diagram of plantwide coal to SNG is illustrated in Figure 1. After being
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crushed and screened, coal with the diameter from 6 to 50 mm is fed into the
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gasification unit. The coal reacts with gasifying agents, which are steam and pure
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oxygen that comes from ASU, converting into crude syngas in BGL gasifier. The
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crude syngas is then washed through scrubber to remove most impurities. The crude
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to WGS unit that used to adjust the syngas composition. In WGS unit, CO reacts with
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H2O to form H2 and CO2, adjusting H2/CO ratio to the range 3.1 to 3.3 required by
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methanation unit. The reaction heat is recovered by generating steam and preheating
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boiler feed water (BFW). After cooled by cooling water, the cooled syngas is then
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sent to AGR unit. The solvent methanol can absorb H2S and CO2 greatly at low
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temperature and is characteristic of high selectivity towards sour components. The
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purified syngas subsequently goes to methanation unit to form SNG. Methanation is a
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strongly exothermic process. Waste heat boiler is used to recover reaction heat by
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producing high pressure (HP) steam. The SNG product is finally sent to pipeline. In
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this work, a 4 billion Nm3/a coal to SNG process is modeled and analyzed. Aspen
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PlusTM (version 8.4) is used for modeling and simulation.
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2.1. Air Separation Unit
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ASU is used to supply high-purity oxygen for the gasification unit. Cryogenic air
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separation is the technology generally used in the ASU. At present, there are mainly
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two kinds of cryogenic air separation technical process, which are internal
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compression process and external compression process.9 In terms of energy
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consumption, the two processes are roughly equivalent. But the internal compression
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has a reputation for more reliability. This is very significant for the gasification unit,
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because oxygen production is at the beginning of the process, and the lack of oxygen
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would bring the whole downstream facility to a standstill. Therefore, an internal
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compression cryogenic air separation is adopted in this project. The capacity of ASU
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is determined by the amount of O2 required by gasification.
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The simulation flowsheet for ASU is shown in Figure 2. Air firstly passes filter
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to remove dust and other solid impurities and is then compressed to 0.62 MPa. After
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cooled by cooling tower, the air is purified by molecular sieve absorbers to remove
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impurities such as steam and carbon dioxide. The purified air drops to -170°C after
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cooled by the first heat exchanger MHEX1 (multi-stream heat exchanger, MheatX
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module) and the second heat exchanger MHEX2 (MheatX module). The partially
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liquefied air is fed into the bottom of the HPC, the oxygen rich liquefied air that
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comes from bottom of HPC returns to the top of the LPC after throttling. Oxygen with
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purity more than 99% and nitrogen can be obtained simultaneously through
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distillation.
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Components involved in this process include O2, N2 and Ar, etc. PENG-ROB is
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selected as the property method.10,11
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2.2. Coal Gasification Unit
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BGL gasifier is a slagging moving bed gasifier which developed based on Lurgi
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gasifier. The main difference of them is that BGL gasifier has a higher reaction
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temperature through facility improvement. Compared with the Lurgi gasifier, the BGL
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gasifier has lower steam consumption and higher CO and H2 yields. At the same time,
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the CO2 content of the crude syngas is lower. The means the BGL gasifier has a
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higher carbon utilization efficiency.
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2.2.1. Description of the BGL Coal Gasifier
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BGL coal gasifier is a countercurrent, pressurized moving bed gaisifer. The
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schematic diagram of the BGL gasifier is shown in Figure 3. The whole gasification
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process in the gasifier bed can be divided into five individual zones: drying, pyrolysis,
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gasification, combustion and overall heat recovery.
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The coal with diameter from 6 to 50 mm is fed into gasifer by coal locks from
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the top. The gasifying agents are injected at the bottom of the bed. The first zone is
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drying zone, in which coal contacts with the gas coming from bottom, and the
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moisture in the coal is released into the gaseous phase. Then the dried coal moves
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downward under gravity to pyrolysis zone, and decompose its volatiles consisting of a
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mixture of char, coal gas and liquids. After the pyrolysis zone, the formed char flows
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combustion zone. Finally, the remaining char reaches the combustion zone and reacts
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with O2, and provides heat for the gasification zone and overall heat recovery.12
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2.2.2. Modeling of the BGL Gasifier
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In this work, the BGL gasifier is modeled by Aspen Plus. Additionally,
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FORTRAN is used to model the chemical kinetic reactions. Shengli lignite is used as
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feedstock with its proximate and elementary analyses listed in Table 1.
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Coal and char can be treated as unconventional componets in the Aspen Physical
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Databas. The built-in models, HCOALGEN and DCOALIGT, are used to calcculate
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their enthalpy. Conventional components involved in gasification include C, CO, CO2,
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CH4, H2, H2O, O2, N2, H2S, S, etc. RK-Soave physical method is selected to calculate
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their properties.13,14 Figure 4 shows the simulation flowsheet for BGL gasification.
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According to the actual processes, five Aspen Plus blocks are used to simulate the
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gasifier: drying zone, pyrolysis zone, gasification zone, combustion zone and overall
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heat recovery.
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In the drying zone, heat exchange is conducted between syngas and the dropping coal, for volatilizing external moisture. The coal drying reaction is shown in Eq. (1).
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Coal (wet) →H2O + Coal (dry) (R1)
(1)
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The amount of vaporized water is determined based on the water content in the
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proximate analysis of coal. A RYield block is used to simulate dehydration of coal in
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the drying zone. The reaction temperature and pressure are set to 240°C and 3.8 MPa.
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The RYield reactor is controlled by a piece of FORTRAN codes.
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The dried coal is then sent to pyrolysis zone. Coal pyrolysis is to break the coal
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to produce CO, CO2, H2, H2O, CH4, H2S, tar, and char, as shown in Eq.(2). In the
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model, tar is represented by C6H6.
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Coal →H2 + CO + CO2 + CH4+ H2O + H2S+N2 + C6H6 +Char (R2)
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Coal pyrolysis is a complex physiochemical process.15-16 The pyrolysis model
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based on experiment to predict the product distribution is more accurate and
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convenient for practical application. The pyrolysis experiments are mainly conducted
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in normal atmosphere, while the operating pressure of the gasifier is about 4 MPa,
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which could not be directly applied to the calculation of the pyrolysis process.17 In this
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work, the effect of pressure difference between industrial condition and experimental
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condition on the distribution of pyrolysis products is corrected by a piece of Fortran
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code embedded in the model.
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In the process of gasification and combustion, homogeneous reactions (gas-gas
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reactions) and heterogeneous (gas-solid reactions) reactions are considered and list in
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Table 2. Due to the different mechanisms of the two kinds of reactions, different
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kinetic expression forms are needed. Table 3 lists the corresponding macro-reaction
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kinetics and reaction rate constants for these reactions.18,19 In the gasification zone, a
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series of CSTRs with external Fortran code are used to simulate the gasification
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process. A RGibbs block is applied to simulate combustion process for the high
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temperature of this zone.
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The model is validated using data obtainde from industrial scale BGL gasifier.
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Table 4 lists the opreational conditions and configuration parameters of BGL gasifier,
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including operating pressure, gasifier height, and gasifier diameter. Based on above
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input conditions, we get the results of the models. Table 5 shows the simulation data
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and industrial data, and their relative error is less than 3%, showing that the
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simulation data and industrial data are in good agreement.20
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2.2.3. Key Operating Parameters Analysis
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The mass ratios of oxygen to coal and steam to coal are two key parameters for
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the gasifier. They directly affect the composition and flow rate of the syngas, and
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further affect the technical and economic performance of the entire downstream
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processes. This paper focuses on the analysis of the effects of oxygen to coal and
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steam to coal ratios on the proportion of effective gas (CO+H2) in the syngas and the
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flow rate of crude syngas, aiming to obtain the reasonable operating parameters.
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When studying the effects of the mass ratio of steam to coal, the other operating
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parameters are fixed while varying the ratio between 0.3 and 0.4. The role of steam is
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to regulate H2 content in the crude syngas, but the excessive amount of steam is a
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waste, and it will also cause adverse effects on the composition of the syngas. Effects
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of steam/coal mass ratio are presented in Figure 5. It is observed that the proportion of
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effective gas decreases with the increase of the mass ratio of steam to coal, but what is
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opposite is the case for the flow rate of the crude syngas.
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Other operating parameters are also fixed when studying the effects of mass ratio
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of oxygen to coal. The variation range of the ratio of oxygen to coal is 0.25~0.50. The
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results are illustrated in Figure 6. The proportion of CO increases while the proportion
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of H2 decreases with the increase of mass ratio of oxygen to coal. For the proportion
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of effective gas in the syngas, there is a maximum value of 75%. The corresponding
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ratio of oxygen to coal is 0.36, and the ratio of steam to coal is 0.35, which are the
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optimum operating parameters.
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2.3. Water Gas Shift and Syngas Cooling Unit
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The H2/CO mole ratio of syngas from BGL gasification is usually between 0.5 to
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0.6. It should be adjusted to higher than, or equal to, the stoichiometric ratio of
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methanation. In this work, the ratio of H2/CO is adjusted to the range 3.1 to 3.3.21
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Figure 7 shows the simulation flowsheet for WGS unit. The crude syngas firstly
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goes into a separator to remove liquid components (Sep module). The gaseous
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components from the separator are preheated to 260°C by the gas-to-gas HEX (HeatX
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module), and is fed into two parallel pre-WGS reactors (RPlug module). Afterthat, 95%
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of the syngas enters two main WGS reactors in series for further shift reaction. The
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outlet steam of the WGS reactor 1 is above 300°C, and goes through the gas-to-gas
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HEX to preheat the feed gas. The syngas out of WGS reactor 2 is 458°C and is cooled
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to 176°C by a LP waste heat boiler (Heater module).After heat recovery, the shifted
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gas is mixed with unshifted gas to reach the demand H2/CO ratio to 3.1-3.3. The
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mixed steam is cooled to 40°C by a cooler(Heater module) using circulating cooling
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water and is sent to the AGR unit. Several RPlug blocks in Aspen Plus are used to
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simulate the pre-WGS reactors and the two main WGS reactors. In these reactors, the
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shift reaction rate is controlled by the embedded Fortran subroutine. The components
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of the WGS process are mainly small-molecule hydrocarbons. RKS is therefore
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selected as the physical method. In addition, STEAM-TA physical method is adopted
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to calculate the properties of circulating cooling water and steam.22The simulation
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data and industrial data are presented in Table 6. It is seen that the model predictions
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are in good agreement with the practical data.21
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2.4. Acid Gas Removal Unit
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The shifted gas from the WGS unit contains amount of CO2 and H2S. CO2 will
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affect methanation reactions, while H2S will lead to the poisoning of methanation
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catalyst. Therefore, it is need to remove CO2 and H2S in AGR unit before the
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methanation unit. Rectisol is widely used in the coal-based SNG projects. The partial
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pressure of CO2 and H2S in the syngas is high. Therefore, it is convenient to remove
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CO2 and H2S by physical absorption method.23,24
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It is well known that conventional gas components (H2, N2, CO, Ar, CH4, CO2,
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H2S, CH3OH and H2O) of Rectisol process will deviate from the ideal state at high
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pressure and low temperature. This would influence the accuracy of simulation results.
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Tremendous efforts have been made to modify physical properties of these substances
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in Aspen plus, especially at the condition of low temperature (-20°C to -50°C). PSRK
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and PC-SAFT, the physical methods, are commonly used for Rectisol process at
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present. The PC-SAFT method with modified binary interaction parameters is
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selected as the physical method for Rectisol process.25
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The simulation flowsheet for AGR unit is presented in Figure 8. The shifted
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syngas (3.75MPa, 40°C) that comes from the WGS unit is mixed with little amount of
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methanol. The purpose of additive methanol is to prevent the pipelines and equipment
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being freezed. The mixture is then cooled to -10°C by a multi-stream cooler. The
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syngas is further cooled down to -20°C by a condenser and fed to the bottom of the
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absorption column. H2S and COS are dissolved in the cooling methanol solvent to a
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larger extent than CO2. The absorption column is divided into two sections: the upper
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section is used for CO2 removal, while the lower for H2S and COS absorption by
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rich-CO2 solvent. The purified syngas goes out from the top of absorption column,
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which is available for methanation. The CO2 and H2S content in the purified syngas
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are reduced by 1.5% and 0.1 ppm. There are two rich solvent streams exiting from
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absorption column: H2S/CO2 rich methanol from the column bottom and the spare
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CO2 rich solvent drawn out from the division of the two sections. The CO2 rich
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solvent from the bottom of the upper section flashes at 1.1MPa in the flash drum for
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CO and H2 recovery. The rich solvent is split into two streams: one goes to CO2
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desorption column for CO2 removal and the other goes to H2S enrichment column to
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absorb H2S in the ascending gas in the column. The CO2 retrieved from the top of
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desorption column is sent to carbon capture and storage (CCS) system. The H2S/CO2
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rich solvent from the bottom of the absorption column flashes at 1.2MPa in the flash
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drum to recycle CO and H2 back to the raw syngas. The rich solvent is then sent to the
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bottom of CO2 desorption column to desorb CO2 by reducing pressure and then goes
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to H2S enrichment column to desorb CO2 by N2 stripping. The H2S rich solvent comes
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from H2S enrichment column is regenerated by distillation, and the H2S rich gas is
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sent to Claus process for sulfur recovery, which is not considered in this work. The
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lean methanol is dehydrated in methanol/water separation column to obtain anhydrous
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methanol and is cooled through multistage cooling to -50°C for circulation. Table 7
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shows comparison between the simulation data of the purified syngas and the
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practical industrial data. It shows that the simulation data matches the industrial data
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very well.7
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2.5. Methanation Unit
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In methanation unit, the sulfur-free syngas converted to methane-rich gas. The
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main reactions taking place in methanation reactors can be described as Eq.(3), Eq.(4)
286
and Eq.(5).
287
CO + 3H2 → CH4 + H2O
∆H = - 206 kJ/mol (R10)
（3）
288
CO2 + 4H2 → CH4 + 2H2O
∆H = - 165 kJ/mol (R11)
（4）
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CO + H2O → CO2 +H2
∆H = - 41 kJ/mol (R12)
（5）
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As can be seen from the heats of reactions, they are all highly exothermic
291
reactions. Therefore, the reaction heat should be removed in time to ensure the
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activity of the catalyst. CO methanation reaction (R10) and water gas shift reaction
293
(R12) are considered as two independent reactions. The kinetic models and
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corresponding parameters of these two reactions are listed in Table 8 and Table 9.27- 29
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The methanation flowsheet is illustrated in Figure 9. The purified syngas from
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the AGR unit is divided into two streams: 40% syngas is fed into the first methanation
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reactor (RPlug module), while the rest is fed into the second methanation reactor
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(RPlug module). The feed gas to first methanation reactor is mixed with the
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circulating gas that comes from the outlet of second methanation reactor to dilute the
300
concentration of CO and H2. By doing this, the temperature of the reactor can be
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controlled in a reasonable range. The heat of the outlet gas from the first two reactors
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is recovered by the waste heat boiler and superheater (Heater module) for generating
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high pressure steam. The cooled gas from the outlet of second methanation reactor is
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also split into two streams: one is for circulation and the other is fed into the third
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methanation reactor (RPlug module) for further methanation, whose outlet gas is
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separated into gas and condensate by a flash vessel. The gas is sent to the fourth
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methanation reactor (RPlug module ), and the outlet gas also passes through a flash
308
vessel to remove the condensate. The methane content should be larger than 96% and
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CO concentration should be less than 100ppm for the final SNG product. The
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reactions rates are controlled by embedded Fortran subroutine. The simulation data is 12 / 48
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311
listed in Table 10, and it is verified by comparing the industrial data.11
312
3. TECHNO-ECONOMIC ANALYSIS
313
3.1. Energy Efficiency Analysis
314
There are two types of energy efficiencies: objective efficiency and general
315
efficiency.30,31 The former is defined as the ratio of the energy of product to the total
316
input energy, while the latter is the ratio of the energy of product and byproduct to the
317
total input energy, as formulated as Eqs. (6) and (7).
318
feedstock fuel η obj = ECH / ( Ecoal + Ecoal + Eutilities )
(6)
319
feedstock fuel η gen = ( ECH + Ebyproduct ) / ( Ecoal + Ecoal + Eutilities )
(7)
4
4
320
where the energy of byproduct Ebyproduct includes the energy of tar, naphtha, phenol,
321
feedstock fuel and etc. The total input energy comprises of feedstock coal Ecoal , fuel coal Ecoal ,
322
and utilities energy Eutilities. The main input-output balance of the Lurgi process and
323
the BGL process are shown in Table 11.
324
According to Eqs. (6) and (7) and Table 11 data, the energy efficiencies of the
325
Lurgi process and the BGL process are calculated and the results are shown in Figure
326
10. It is seen that the objective energy efficiencies of the Lurgi process and the BGL
327
process are 48% and 56%. The BGL process’s energy efficiency increases 8%
328
comparing to that of the Lurgi process. This is mainly because the steam and oxygen
329
consumption of BGL gasifier is only about 12% and 82% respectively to that of Lurgi
330
gasifier. Thus, the energy efficiency of the BGL process has been improved. However,
331
the general energy efficiency of the BGL process is only 3% higher than that of the
332
Lurgi process if considering byproduct in calculation of the energy efficiency. This is
333
because that the Lurgi process can generate more byproducts including tar, naphtha,
334
and phenol.
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335
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3.2. Economic Performance Analysis
336
Total capital investment and production cost of the BGL process are analyzed.
337
They would be compared with that of Lurgi process. These two processes are
338
assumed as 4 billion Nm3/y in this paper.
339
3.2.1. Total Capital Investment
340
Total capital investment is composed of fixed capital investment and working
341
capital investment. Fixed capital investment is used to purchase equipment,
342
instruments, piping, land, etc. Working capital investment is used to maintain
343
operation. It is composed of inventory of raw materials, finished products,
344
semi-finished products, and taxes, etc.
345
Total capital investment usually depends on the main equipment investment. The
346
other components of the total capital investment could be estimated according to their
347
ratios to the equipment investment. Orhan listed a detailed ratio factors, which is
348
shown in Table 12.32,33
349
Aspen Process Economic Analyzer is used to estimate the equipment investment.
350
The total investment cost updates to 2015-USD TCI2015 is obtained after applying
351
Eq.(8).:
I 2015−Y ) 100
352
TCI 2015 = TCIY (1 +
353
Where: TCI is the total investment cost; Y is the start-up year of reference plant
354
(8)
(Lurgi process); I is the average inflation rate in China in the period 2013-2015.
355
The equipment investment of the BGL process is 0.15 billion USD, the total
356
capital investment of the BGL and the Lurgi processes are calculated, which are 4.31
357
billion USD and 3.85 billion USD, respectively. As shown in Figure 11, it is observed
358
that the total capital investment of the BGL process is 12.1% higher than that of the
359
Lurgi process. This is because that: the BGL process requires more syngas to produce
360
the same amount of SNG owing to lower CH4 content in the crude syngas14. The load
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of gasifier increases and equipment investment also increased accordingly. Due to the
362
lower mole ratio of H2/CO, more syngas is fed into WGS process and leads to the
363
increase of handling load of AGR process and methanation process, which makes the
364
investment increased as a result.
365
3.2.2. Production Cost
366
Production cost is another important economic indicator to evaluate the
367
economic performance of a process. Production cost (PC) is the sum of manufacturing
368
costs, administrative costs, selling costs and others. PC is calculated on the basis of
369
Eq. (9). The manufacturing costs include direct production costs, fixed charges and
370
plant overhead costs. Some assumptions are made as shown in Table 1333. Raw
371
materials prices are average prices in 2015.
372
PC=CR+CU+CO&M+CD+CPOC+CAC+CDSC
(9)
373
Where CR is the raw materials cost; CU is the utilities cost; CO&M is the operating
374
and maintenance cost; CD is the depreciation cost; CPOC is the plant overhead cost;
375
CAC is the administrative cost; CDSC is the distribution and selling cost.
376
The production costs of BGL process and Lurgi process are 0.218 USD/m3 SNG
377
and 0.241 USD/m3 SNG shown in Figure 12. The production cost of BGL process is
378
decreased by 9.5% compared to Lurgi process. The main reasons of decrease are as
379
follow. BGL gasifier consumes less steam, which is only about 20% that of Lurgi
380
process. As a result, the utilities cost of BGL process is lower than that of Lurgi
381
process. Producing 1 m3 SNG, BGL process consumes 2.21kg coal, while Lurgi
382
process consumes only 2.13kg coal, so the raw materials cost of BGL process is
383
increased by 3.8%. Depreciation expense is positively related to total capital
384
investment. In the previous analysis, the total capital investment of BGL process is
385
more than that of Lurgi process, the depreciation expense is therefore more than Lurgi
386
process. However, the decrease range of utilities cost is higher than the increase range
387
of raw materials cost and depreciation expense, so that production cost of BGL
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388
process is lower than that of Lurgi process.
389
3.2.3. Return on Investment
390
Return on investment (ROI) is the ratio of total net profit over the total capital
391
investment.35 According to the current market prices of product and byproduct, the
392
incomes of the Lurgi process and the BGL process are calculated. One scenario is
393
only consider the income of product, while the other scenario includes the incomes of
394
product and byproduct. In first case, these two processes have same income 1515
395
M$/y. In another case, the income of the BGL process is 1590 M$/y, while that of the
396
Lurgi process is 1848 M$/y. The project construction period is assumed to be one year,
397
while the production period is 20 years. The total net profits of the BGL and the Lurgi
398
processes are 666/742 M$/y and 455/788 M$/y in two scenarios. Thus, the ROIs of
399
these three processes are 15/17% and 12/20%, as shown in Figure 13.
400
It is clear that the ROI of the BGL process is 3% higher than that of the Lurgi
401
process in first case. This is because the BGL process’s feedstock consumption for 1
402
Nm3 SNG is much lower than that of the Lurgi process, which results in lower
403
production cost. However, the Lurgi process can produce more byproducts. If
404
considering values of these byproducts, the BGL process will not has economic
405
competitive.
406
4. CONCLUSIONS
407
In our work, we put forward a novel BGL gasifier based coal to SNG process.
408
This process includes an air separation unit, a coal gasification unit, a water gas shift
409
unit, an acid gas removal unit, and a methanation unit. These units are simulated
410
based on rigorous kinetics models. Depending on simulation, mass flow and energy
411
flow are obtained. Then, analyses of energy efficiency and economic performance are
412
conducted.
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Results show that the objective energy efficiency of the BGL process is 56%, 8%
414
higher than that of the Lurgi process. This is due to the BGL process consumes 3.6 t
415
coal for producing 1 Nm3 SNG, while the Lurgi process consumes 4.0 t coal.
416
However, the general energy efficiency is only 3% larger than that of the Lurgi
417
process. Because the Lurgi process produces more byproducts including tar, naphtha,
418
and phenol. As smaller feedstock consumption for 1 Nm3 SNG, the BGL process has
419
more economic competitiveness with 3% of return on investment higher than that of
420
the Lurgi process. However, this competitiveness will not existed taking into account
421
of the production of byproducts.
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422
AUTHOR INFORMATION
423
Corresponding Author
424
E-mail:
[email protected]. Tel.: +86-18588887467
425
Notes
426
The authors declare no competing financial interest.
427
ACKNOWLEDGMENTS
428
The authors are grateful for financial support from the China NSF projects
429
(21136003
430
(2014CB744306).
431
NOMENCLATURE
432 433 434 435 436 437 438 439 440
CR CU CO&M CD CPOC CAC CDSC E feedstock Ecoal
441 442 443 444 445 446 447 448 449 450
Page 18 of 48
fuel E coal
Eutilities Ebyproduct
ηobj ηgen
and
21306056)
and
the
National
Basic
Research
raw materials cost (USD/y) utilities cost (USD/y) operating and maintenance cost (USD/y) depreciation cost (USD/y) plant overhead cost (USD/y) administrative cost (USD/y) distribution and selling cost (USD/y) chemical reaction activation energy (kJ·mol-1) energy of feedstock (MW) energy of fuel (MW) utilities energy (MW) byproduct energy including the energy of tar, naphtha, phenol, etc. (MW) objective energy efficiency (%) general energy efficiency (%)
Abbreviations AGR acid gas removal BFW boiler feed water 18 / 48
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Program
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Energy & Fuels
451 452 453 454 455 456 457 458 459 460
HP LP MP WGS RYield RCSTR RGibbs Rstoic Sep TPC
461
References
462 463 464 465 466 467 468 469 470 471 472 473 474 475 476 477 478 479 480 481 482 483 484 485 486 487 488 489
1.
high pressure steam low pressure steam medium pressure steam water gas shift yield reactor continuous stirred tank reactor gibbs reactor stoichiometric reactor separator total capital cost
Wang, J.; Jiang, H.; Zhou, Q.; Wu, J.; Qin, S. China’s natural gas production and consumption analysis based on the multicycle Hubbert model and rolling Grey model. Renewable and Sustainable Energy Reviews 2016, 53, 1149-1167.
2.
International Energy Agency (IEA). World Energy Outlook 2015; IEA: Pairs, France, 2015.
3.
Kong, Z.; Dong, X.; Liu, G. Coal-based synthetic natural gas vs. imported natural gas in China: a net energy perspective. Journal of Cleaner Production 2016, 131, 690–701.
4.
Li, H.; Yang, S.; Zhang, J.; Kraslawski, A.; Qian, Y. Analysis of rationality of coal-based synthetic natural gas (SNG) production in China. Energy Policy 2014, 71 (3), 180-188.
5.
Ding, Y.; Han, W.; Chai, Q.; Yang, S.; Shen, W. Coal-based synthetic natural gas (SNG): A solution to China’s energy security and CO2 reduction? Energy Policy 2013, 55, 445-453.
6.
Yang, C. J.; Jackson, R. B. China's synthetic natural gas revolution. Nature Climate Change 2013, 3 (10), 852-854.
7.
Yu, B. Y.; Chien, I. Design and economic evaluation of a coal-to-synthetic natural gas process. Ind. Eng. Chem. Res. 2015, 54 (8).
8.
Cooke, B. H.; Taylor, M. R. The environmental benefit of coal gasification using the BGL gasifier. Fuel 1993, 72 (3), 305-314.
9.
Lin, Z. W. Comparison and option of internal-compression and external-compression process for air separation unit. Cryogenic Technology 2007.
10. Ham, L. V. V. D.; Kjelstrup, S. Improving the heat integration of distillation columns in a cryogenic air separation unit. Ind. Eng. Chem. Res 2011, 50 (15), 9324-9338. 11. Fu, Q.; Zhu, L.; Chen, X. Complete equation-oriented approach for process analysis and optimization of a cryogenic air separation unit. Ind. Eng. Chem. Res 2015, 54 (48). 12. Jian, X.; Yong, Y.; Li, Y. W. Recent development in converting coal to clean fuels in China. Fuel 2015, 152, 122-130. 13. He, C.; Feng, X.; Chu, K. H. Process modeling and thermodynamic analysis of Lurgi fixed-bed coal gasifier in an SNG plant. Applied Energy 2013, 111 (11), 742-757. 14. Nikoo, M. B.; Mahinpey, N. Simulation of biomass gasification in fluidized bed reactor using ASPEN PLUS. Biomass & Bioenergy 2008, 32 (12), 1245-1254.
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490 491 492 493 494 495 496 497 498 499 500 501 502 503 504 505 506 507 508 509 510 511 512 513 514 515 516 517 518 519 520 521 522 523 524 525 526 527 528 529 530 531
15. Anthony, D. B.; Howard, J. B. Coal devolatilization and hydrogastification. AICHE Journal 1976, 22 (4), 625-656. 16. Chen, H.; Luo, Z.; Yang, H.; Ju, F. Zhang, S. Pressurized Pyrolysis and Gasification of Chinese Typical Coal Samples. Energy & Fuels 2008, 22 (2), 1136-1141. 17. Yang, H.; Chen, H.; Ju, F.; Yan, R.; Zhang, S. Influence of pressure on coal pyrolysis and char gasification. Energy & Fuels 2007, 21 (6), 3165-3170. 18. Krishnudu, T.; Madhusudhan, B. Reddy, S. N. Studies in a moving bed pressure gasifier. Ind. Eng. Chem. Res. 1989, 28:4. 19. Aspen Technology. Model for moving bed coal gasifier. 2010 20. Zheng, L.; Furinsky, E. Comparison of Shell, Texaco, BGL and KRW gasifiers as part of IGCC plant computer simulations. Energy Conversion & Management 2005, 46 (11-12), 1767-1779. 21. Li, S.; Ji, X.; Zhang, X.; Lin, G.; Jin, H. Coal to SNG: Technical progress, modeling and system optimization through exergy analysis. Applied Energy 2014, 136, 98-109. 22. Bustamante, F.; Enick, R. M.; Cugini, A. V.; Killmeyer, R. P.; Howard, B. H.; Rothenberger, K. S.; et al. High-temperature kinetics of the homogeneous reverse water–gas shift reaction. AICHE Journal 2004, 50 (50), 1028-1041. 23. Yang, S.; Qian, Y.; Yang, S. Development of a Full CO2 Capture process based on the rectisol wash technology. Ind. Eng. Chem. Res 2016. 24. Li, S.; Smith, R. Rectisol wash process simulation and analysis. Journal of Cleaner Production 2013, 39 (1), 321-328. 25. Koytsoumpa, E. I.; Atsonios, K.; Panopoulos, K. D.; Karellas, S.; Kakaras, E.; Karl, J. Modelling and assessment of acid gas removal processes in coal-derived SNG production. Applied Thermal Engineering 2015, 74 (1), 128-135. 26. Salazar, J. M.; Diwekar, U. M.; Zitney, S. E. Rigorous-simulation pinch-technology refined approach for process synthesis of the water–gas shift reaction system in an IGCC process with carbon capture. Computers & Chemical Engineering 2011, 35 (9), 1863-1875. 27. Kopyscinski, J.; Schildhauer, T. J.; Biollaz, S. M. A. Fluidized-bed methanation: interaction between kinetics and mass transfer. Ind. Eng. Chem. Res 2010, 50 (5), 2781-2790. 28. Xu, J.; Froment, G. F. Methane steam reforming, methanation and water-gas shift: I. Intrinsic kinetics. AICHE Journal 1989, 35 (1), 88–96 29. Kopyscinski, J.; Schildhauer, T. J.; Biollaz, S. M. A. Production of synthetic natural gas (SNG) from coal and dry biomass – A technology review from 1950 to 2009. Fuel 2010, 89 (8), 1763-1783. 30. Kao, Y. L.; Lee, P. H.; Tseng, Y. T.; Chien, I. L.; Ward, J. D. Design, control and comparison of fixed-bed methanation reactor systems for the production of substitute natural gas. Journal of the Taiwan Institute of Chemical Engineers 2014, 45 (5), 2346-2357. 31. He, C.; You, F., Shale gas processing integrated with ethylene production: novel process designs, exergy analysis, and techno-economic analysis. Ind. Eng. Chem. Res 2014, 53 (28), 11442-11459. 32. Qian, Y.; Man, Y.; Peng, L.; Zhou, H., Integrated process of coke-oven gas tri-reforming and coal gasification to methanol with high carbon utilization and energy efficiency. Ind. Eng.
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Chem. Res 2015, 54 (9), 2519-2525. 33. Man, Y.; Yang, S.; Qian, Y. Integrated process for synthetic natural gas production from coal and coke-oven gas with high energy efficiency and low emission. Energy Conversion & Management 2016, 117, 162-170. 34. Zhou, H.; Yang, S.; Xiao, H.; Yang, Q.; Qian, Y.; Gao, L. Modeling and techno-economic analysis of shale-to-liquid and coal-to-liquid fuels processes. Energy 2016, 109, 201-210. 35. Cho, H. J.; Kim, J. K.; Cho, H. J.; Yeo, Y. K. Techno-economic study of a biodiesel production from palm fatty acid distillate. Ind. Eng. Chem. Res 2013, 52 (1), 462-468.
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541
Caption of Figures and Tables
542
Figures
543
Figure 1. Flow diagram for plantwide coal to SNG process.
544
Figure. 2. Simulation flowsheet for ASU.
545
Figure 3. Schematic diagram of the BGL gasifier.
546
Figure 4. Simulation flowsheet for BGL gasification
547
Figure 5. Effects of steam/coal mass ratio.
548
Figure 6. Effects of O2/coal mass ratio.
549
Figure 7. Simulation flowsheet for WGS unit.
550
Figure. 8. Simulation flowsheet for AGR unit.
551
Figure 9. Simulation flowsheet for methanation unit.
552
Figure 10. Energy efficiency of BGL process and Lurgi process.
553
Figure 11. Total capital investments of the BGL and Lurgi processes.
554
Figure 12. Production costs of the BGL and Lurgi processes.
555
Figure 13. Return on investment: (a) income of product (b) incomes of product and byproduct.
556 557
Tables
558
Table 1. Properties of Shengli Coal.
559
Table 2. Reactions Considered in the Gasification and Combustion Zone.
560
Table 3. Reaction Kinetics for the Gasification and Combustion Processes.
561
Table 4. Operational Conditions and Configuration Parameters of BGL Gaisifier.
562
Table 5. Comparison of Simulation Data and the Industrial Data of the Gasification.
563
Table 6. Comparison of Simulation Data and the Industrial Data of the WGS Unit.
564
Table 7. Comparison of Simulation Data and the Industrial Data of the Rectisol.
565
Table 8. Reaction Kinetics for Methanation Unit.
566
Table 9. Kinetics Parameter Value.
567
Table 10. Comparison of Simulation Data and the Industrial Data of the Mathanation Unit.
568
Table 11. The Input-output Balance of the BGL Process and the Lurgi Process.
569
Table 12. Percentage of Components in Investment.
570
Table 13. Estimation Coefficient of Production Cost.
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Energy & Fuels
Coal
Acid Gas WGS unit
LP steam Air ASU
O2
Gasifier
BFW
MP steam Steam BFW
AGR unit
Fired Boiler ASH
Waste water
SNG Product
571
Methanation unit
572 573
Figure 1. Flow diagram for plantwide coal to SNG process.
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Page 24 of 48
30℃ 30℃ 0.55MpaO2 N2 0.101Mpa
-178℃ 0.102Mpa
17℃ 0.62Mpa MHEX1 Molecular Sieve Absorbers LPC Air 23.5℃ 0.089Mpa
Filter
-175℃ 0.55Mpa
17℃ 0.62Mpa
-116℃ 3.9Mpa MHEX2
23.2℃ 0.088Mpa
Spray Cooler
HPC -171℃ 0.57Mpa
T-1 -173℃ 0.58Mpa
Air Compressor
574 575
105℃ 0.62Mpa
Compressor
Figure. 2. Simulation flowsheet for ASU.
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T-2
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Energy & Fuels
Coal Lock hopper
Coal
Syngas
Aqueous liquor
Drying zone Distributor
Pyrolysis zone Refractory lining Aqueous liquor
Gasification zone
Steam
Steam and oxygen
O2 + H2O Combustion zone Slag quench chamber
Circulating quench water
Ash
Slag lock hopper
576 577
Figure 3. Schematic diagram of the BGL gasifier.
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Syngas
Page 26 of 48
180℃ 4.0MPa
150℃ 0.5MPa CW
161℃ 0.5MPa
150℃ 5.0MPa
LPS
CW
435℃ MPS 4.5MPa Water jacket
Heat recovery
40.31t/h 15-50mm
110℃ O2 4.5MPa
Coal
MPS 435℃ 4.5MPa ASH
SEP-1
578 579
Drying
SEP-2 Gasification
Pyrolysis
Combustion
Figure 4. Simulation flowsheet for BGL gasification .
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200℃ 4.0MPa
Page 27 of 48
90
800
89
760
88
720
87
680
CO+H2
86 0.30
0.32
0.34
0.36
0.38
Steam/coal(kg·kg-1）
580 581
Figure 5. Effects of steam/coal mass ratio.
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640 0.40
Crude syngas/103m3·h-1
Crude syngas
m(CO+H2)/%
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60
Energy & Fuels
Energy & Fuels
88.5
80
88.0 60
CO+H2
40
583
87.0
H2
86.5
20
CO2 0.25
582
87.5
0.30
0.35
0.40
0.45
O2/coal(kg·kg-1）
Figure 6. Effects of O2/coal mass ratio.
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86.0 0.50
Efficient Synthesis Gas Yield(vol.%)
CO Product Gas Composition (vol%)
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60
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Energy & Fuels
3.82Mpa 260℃
MP steam 4.8Mpa 450℃
3.92Mpa 260℃ 3.82Mpa 240℃
WGS reactor1
Pre-WGS reactor Crude gas from gasification
4.0Mpa 181℃
Gas-togas HEX
Crude gas separator
4.0Mpa 181℃
3.86Mpa 325℃ 3.9Mpa 294℃
WGS reactor2
3.76Mpa 458℃
Gas liquor MP steam
BFW LP waster heat boiler
3.75Mpa 282℃
Cooler
584 585
Figure 7. Simulation flowsheet for WGS unit.
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3.74Mpa 40℃
Energy & Fuels
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Nitrogen
H2S Fraction
Wash Water Tail Gas 30℃ 0.15Mpa
Syngas From WGS
-60℃ 0.2Mpa
-51℃ 0.3Mpa -48℃ 5.5Mpa
Absorption column Refr
-16℃ 5.55Mpa
-19℃ 5.55Mpa
76℃ 0.6Mpa
Cooling
Flash Regenerator
40℃ 0.3Mpa
-29℃ 0.3Mpa
MeOH Injection
32℃ 0.12Mpa
98℃ 0.3Mpa
-50℃ 5.5Mpa
40℃ 5.6Mpa
34℃ 0.18Mpa
32℃ 0.3Mpa
-57℃ 0.2Mpa
Steam
-53℃ 0.2Mpa
-12℃ 5.6Mpa -31℃ 3.3Mpa
-12℃ 5.6Mpa
-38℃ 0.3Mpa
84℃ 0.3Mpa
96℃ 0.3Mpa 60℃ 0.4Mpa
Steam 37℃ 0.15Mpa 80℃ 0.4Mpa
-37℃ 3.3Mpa
CO2 Product Syngas To Methanation
586 587
30℃ 0.3Mpa -12℃ 5.6Mpa 40℃ 0.1Mpa
30℃ 5.5Mpa
Figure. 8. Simulation flowsheet for AGR unit.
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Waste Water
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Energy & Fuels
250 ℃ 3.34 Mpa
280 ℃ 3.38 Mpa 320 ℃ 3.5 Mpa
First methanation Reactor
Syngas from Rectisol
Second methanation Reactor
295 ℃ 3.5 Mpa
HEX4
Flash1 HEX3
Waste heat bolier 2 Superheater
SNG
449 ℃ 3.35 Mpa
Waste heat bolier 1
589
40 ℃ 3.33 Mpa
Compressor 620 ℃ 3.44 Mpa
588
Forth methanation Reactor
Third methanation Reactor
320 ℃ 3.45 Mpa
HEX2
40 ℃ 3.33 Mpa
70 ℃ 3.35 Mpa
Flash2
Process condensate
HEX1
Figure 9. Simulation flowsheet for methanation unit.
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Flash3
Process condensate
Energy & Fuels
60 Objective efficiency(%)
56
55 50
48
45 40 35 30 BGL process
590 591
Lurgi process
(a) 70
General efficiency（%）
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60
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60
58 55
50 40 30 20
592 593 594
BGL process
Lurgi process
(b) Figure 10. Energy efficiency of BGL process and Lurgi process.
595
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6
Equipment and installation Instrument,piping and electrical Building and land Engineering and supervision Construction and contractors fee Working capital and contingency
5
4 109 USD
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60
Energy & Fuels
3
2
1
0
596 597
BGL process
Lurgi process
Figure 11. Total capital investments of the BGL and Lurgi processes.
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Energy & Fuels
0.3 Raw material Utilities Depreciation expense Others
0.2 USD/m3 SNG
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60
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0.1
0
598 599
BGL process
Lurgi process
Figure 12. Production costs of the BGL and Lurgi processes.
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Return on investment (%)
20
16
15 12
12
8
4
0
BGL process
600 601
Lurgi process
(a) 25 20 Return on investment(%)
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60
Energy & Fuels
20 17 15 10 5 0
602 603 604
BGL process
Lurgi process
(b) Figure 13. Return on investment: (a) income of product (b) incomes of product and byproduct.
605 606
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607
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Table 1. Properties of Shengli Coal. Properties
Value
Proximate analysis, wt% Moisture, wet
17.81
Volatile matter, dry
35.88
Fixed carbon, dry
49.5
Ash, dry
14.62
Ultimate analysis, wt% Ash, dry
14.62
H, dry
3.89
Cl, dry
0
O, dry
15.3
C, dry
63.55
N, dry
0.74
S, dry
1.9
608
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Energy & Fuels
609
Table 2. Reactions Considered in the Gasification and Combustion Zone. Reaction No.
Reaction heat,
Reaction
MJ mol-1
R3
C+H2O → CO + H2
R4
C + 2H2 → CH4
R5
C + CO2 → 2CO
R6
C+
R7
CH4 +H2O → CO + 3H2
R8
CO + H2O → CO2 +H2
R9
H2 + 0.5O2 → H2O
1
λ
O2 → (
2
λ
+131 -75 +172
− 1) CO 2 + ( 1 −
1
λ
) CO, λ ∈ [1, 2 ]
ㅡ +206 -41 -242
610
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Energy & Fuels
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611
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Table 3. Reaction Kinetics for the Gasification and Combustion Processes. Reaction
Reaction rate
Rate constant
Units
No. -5051
PH* 2 O =
exp(17.29-
5051 )CC (PH2 -P*H2 ） T
P*H2 = ቈ
expቀ-13.43-
3000exp(
R4
exp（-7.087-
R5
3000exp(
R6
T·dparticle 1 1 ቈ ൬ -1൰ (PO2 ) + 1.244ሺT/1800ሻ1.75 kdiff ·ε2.5 Y b
R7
312exp(-
R8
2.96×10-4 ·T2exp(-
R9
8.83×108 exp ൬-
T
PCO PH
)CC (PH2 O -PH* 2 O ）
R3
-5051 )CC (PCO2 -P*CO2 ） T
2 16330 ) T
mol（cm3s）-1 0.5
PCH
4 10100 ቁ T
mol（cm3s）-1
_
mol（cm3s）-1
_
mol（cm3s）-1
-1
30000 ) 1.987T
PH* 2 O = CCO2 ·CH2 4895 )(CCO ·CH2 O ) T kwgs
12003 2 CCO2 ൰ ൫CH2 ൯ T CCO
PCO expሺ20.92-20280/Tሻ
kwgs =exp(-4.3+
_
612
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4577.8 ) T
s-1
mol（m3s）-1
mol（m3s）-1
Page 39 of 48
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Energy & Fuels
613
Table 4. Operational Conditions and Configuration Parameters of BGL Gaisifier. Parameter
Value
Operating pressure/MPa
3.8
Height/m
12.5
Diameter/m
3.6
614
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615
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Table 5. Comparison of Simulation Data and the Industrial Data of the Gasification. Items
Mole fraction/%（dry basis） CO
H2
CO2
CH4
N2+Ar
H2 S
C2+
NH3
Simulation data
56.99
25.41
8.28
7.63
0.23
0.46
0.98
0.020
Industrial data
57.02
25.39
8.29
7.60
0.26
0.43
0.99
0.029
616
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Energy & Fuels
617
Table 6. Comparison of Simulation Data and the Industrial Data of the WGS Unit. Items
Mole fraction/%（dry basis） CO
H2
CO2
CH4
N2+Ar
H2 S
C2+
Simulation data
12.99
46.23
33.57
5.88
0.19
0.34
0.80
Industrial data
13.76
45.96
33.59
5.51
0.12
0.30
0.76
618
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619
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Table 7. Comparison of Simulation Data and the Industrial Data of the Rectisol. Items
Mole fraction/%（dry basis） CO
H2
CO2
CH4
C2+
Simulation data
21.66
67.10
1.39
8.42
1.13
Industrial data
20.49
68.43
1.47
8.20
1.31
620
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Energy & Fuels
621
Table 8. Reaction Kinetics for Methanation Unit.
Reaction No. R10
R12
Reaction rate
Rate constant
େ ୌଷ మ ݇ௌெோ ܭୗୖ − େୌర ୌమ ݎ1 = ଶ.ହ ୌమ 1 + ܭେ େ + ܭୌమ ୌమ +ܭେୌర େୌర +ܭୌమ ୌమ ୌିଵమ
େ ୌ େ ୌమ − ܭమ మ ݇ௐீௌ ୋୗ ݎ2 = ୌమ ሺ1 + ܭେ େ + ܭୌమ ୌమ +ܭେୌర େୌర +ܭୌమ ୌమ ୌିଵమ ሻଶ
622
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ିா
݇ = ݇ exp ቀ ୖ ቁ , ݅ = SMR, WGS
ܭ = ܭ exp ൬
−△ܪ ൰,݅ RT
= CO, CHସ , Hଶ , Hଶ O
Energy & Fuels
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60
623
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Table 9. Kinetics Parameter Value. Kinetics parameter
Value
Unit
k10
3.71E +14
kmol·pa1/2/（kg·h）
k20
5.43E - 03
kmol/（kg·h·pa）
EA,1
240
kJ/mol
67.13
kJ/mol
8.23
pa-1
1.77E + 05
1
KCH4
66.50
pa-1
KH20
6.12E - 04
pa-1
-70.65
kJ/mol
△H CH4
88.68
kJ/mol
△HH2
-38.28
kJ/mol
-82.90
kJ/mol
EA,2 KCO
0
KH2O0 0
△HCO
△HCH4
624
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Energy & Fuels
625
Table 10. Comparison of Simulation Data and the Industrial Data of the Mathanation Unit. Item
Mole fraction/%（dry basis） H2
CO2
CH4
N2+Ar
Simulation data
0.30
1.52
97.11
1.07
Industrial data
0.31
1.64
96.46
1.58
626
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627
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Table 11. The Input-output Balance of the BGL Process and the Lurgi Process Items
BGL process
Qinghua Lurgi process
Input Feedstock coal (t/h)
1235.2
Fuel coal (t/h)
204.8
810
953
1892
SNG (kNm3/h)
500
687.5
Tar (t/h)
15.4
52.48
Naphtha (t/h)
1.6
6.64
Phenol (t/h)
5.7
12.8
Cooling water (t/h)
1882
Output
628 629
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Energy & Fuels
630
Table 12. Percentage of Components in Investment. Component
RF (%)
(1) Direct investment (1.1) Equipment
100
(1.2) Installation
8
(1.3) Piping
12
(1.4) Instrumentation and controls
12
(1.5) Electrical
8
(1.6) Land
6
(1.7) Buildings (including services)
20
(2) Indirect investment (2.1) Engineering and supervision
12
(2.2) Construction
22
(2.3) Contractors fees
18
(2.4) Contingency
13
(3) Fixed capital
231
(4) Working capital
45
(5) Total capital
276
631 632
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Table 13. Estimation Coefficient of Production Cost a.
633
Component
Economic assumption
（1）Raw materials cost
Coal 24.1 USD/t b
（2）Utilities cost
H2O 0.3 USD/t; Electricity 0.075 USD/kWh
（3）Operating & maintenance (3.1) Operating labor
1500 labors, 1800 USD/labor/year
(3.2) Direct supervisory & clerical labor
634
10 % of operating labor
(3.3) Maintenance and repairs
3 % of fixed capital investment
(3.4) Operating supplies
1 % of fixed capital investment
(3.5) Laboratory charge
10 % of operating labor
（4）Depreciation
Life period 15 years; salvage value 5%
（5）Plant overhead cost
5% (3.1+3.2+3.3)
（6）Distribution and selling cost
2% of production cost
（7）Production cost
(1) + (2) + (3) + (4) + (5) + (6)
a
The exchange rate is 1 RMB = 0.15 USD; b The coal cost is calculated by factory price.
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