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Small scale ammonia production from biomass: A techno-enviro-economic perspective Pratham Arora, A.F.A Hoadley, Sanjay M. Mahajani, and Anuradda Ganesh Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.5b04937 • Publication Date (Web): 06 May 2016 Downloaded from http://pubs.acs.org on May 17, 2016

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Industrial & Engineering Chemistry Research

Small scale ammonia production from biomass: A techno-enviro-economic perspective

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Pratham Aroraa, b, c, Andrew F.A. Hoadleyb*, Sanjay M. Mahajanid, Anuradda Ganeshc

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a

IITB-Monash Research Academy, Indian Institute of Technology – Bombay, Mumbai, 400076, India. E-mail: [email protected]

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b

Department of Chemical Engineering, Monash University, Clayton, VIC-3168, Australia. Email: [email protected]

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c

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d

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Abstract Ammonia production has traditionally been based on large scale plants. The thrust towards large scale production in order to gain economic advantages has overshadowed the benefits that could be derived out of small scale production plants. Additionally, the ammonia industry consumes a major chunk of global fossil fuels, which also burdens the planet with greenhouse gases. To effectively counter these issues, this study investigates the production of ammonia from biomass. Processes based on biomass plants are usually small scale and are limited by biomass supply. In order to ensure sustainable ammonia production, this study tries to highlight the techno-economic advantages that result from small scale ammonia plants based on biomass feedstock. This paper proposes a new process that takes inputs from a relatively old, natural gas based process (Leading Concept Ammonia®) specifically designed for small scale ammonia manufacture and couples it with a recently developed dual fluidized bed technology for biomass feedstock. Two different flowsheet configurations are simulated rigorously and compared to gain a better understanding of the process. The flowsheets are optimized and energy integration is performed to provide a wider insight. The life cycle assessment calculations that are carried out using ASPEN Plus® simulation results and ecoinvent® databases, predict a CO2 emissions reduction of 54-68 % when compared to conventional ammonia plants.

Department of Energy Science and Eng., Indian Institute of Technology-Bombay, Mumbai, 400076, India. E-mail: [email protected]

Department of Chemical Engineering, Indian Institute of Technology-Bombay, Mumbai, 400076, India. E-mail: [email protected]

Keywords: - Biomass gasification, ammonia, ASPEN Plus, techno-economic analysis, optimization, life cycle assessment. *Corresponding author. Tel.: +61 3 990 53421. E-mail address: [email protected] (Andrew Hoadley).

ACS Paragon Plus Environment

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1. Introduction

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Ammonia is often regarded as a chemical that changed the course of the 20th century. As a

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fertilizer, it has sustained global agriculture, and as an explosive, it has become indispensable to

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the mining industry. Development of high-pressure turbo-compressors in 1960’s facilitated the

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building of high-pressure ammonia units with production capacities greater than 1000 ton per

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day (tpd). These large scale plants steadily dominated the ammonia production worldwide, as

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they could cater to the ever increasing ammonia demands. Due to energy efficiency

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improvements, the energy requirements in these plants have come down from 60 GJ/tNH3 to 28

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GJ/tNH3 in the last 50 years. The process arrangement, however, has remained largely unaltered.1

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This development has overshadowed the advantages of small-scale ammonia production. The

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production of ammonia from a world-scale ammonia plant often exceeds its local consumption

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as a fertilizer or for mining activities. A typical 1000 tpd ammonia plant can meet ammonia

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requirements of many small countries. For landlocked countries that are located at a long

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distance from the ports, the freight cost doubles or even triples the cost of ammonia-based

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products, and thus, limits consumption.2 This problem can be solved by decentralised ammonia

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plants of smaller capacity. Smaller plants tend to take advantage of the skid-mounted “serial”

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installation concept in which the installations, in a nearly completed form, are transported from

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the building site. This counters the absence of necessary industrial infrastructure in a remote

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location.

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utilization and lower engineering costs. Ease of training, and quicker and fewer start-ups/shut-

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downs are other added advantages.2

Additionally, the smaller plants have shorter construction times, higher capacity

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On the flipside, conventional ammonia production continues to have a significant environmental

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footprint by consuming 1.2% of the global primary energy and contributing to 0.93% of

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worldwide Greenhouse Gas (GHG) emissions.3 This stems from the fact that natural gas has

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remained the preferred feedstock due to its low price and wide availability. Other feedstocks,

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such as coal and naphtha, due to high carbon content, release more than double the amount of

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GHG’s, making them even less attractive. This has initiated a quest for alternative ammonia

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production feedstocks among researchers world-wide. Biomass, with its wide availability and

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carbon neutrality, stands as a strong candidate for the replacement of fossil fuels. Many authors

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in the recent past have acknowledged the potential of ammonia production from biomass through

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the gasification route.

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Ahlgren et al.4 did a Life Cycle Assessment of ammonium nitrate fertilizer production from

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biomass. They considered two biomass feedstocks, namely, cereal straw and Salix. In their

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cradle-to-grave analysis of the process, they highlighted the effect of a biomass feedstock on the

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global warming potential, acidification potential and eutrophication potential.

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predicted a reduction of 30 % and 22 % in the global warming impact, for the Salix and straw

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feedstocks, respectively. Gilbert et al.5 did an economic analysis for 1200 tpd ammonia plant

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based on roundwood feedstock. They predicted that the cost for the production of ammonia from

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biomass would be competitive to the present market. They also report a greenhouse gas

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reduction of 65% when compared to conventional natural gas fed ammonia plants. Their

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proposed plant would, however, require 1,121,000 tons per annum (tpa) of biomass feedstock.

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This is greater than the world’s largest biomass power plant. A similar techno-environomic study

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including carbon capture was carried out by Tock et al.6 for a 1187 tpd ammonia plant. They

They have


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predicted a need for carbon tax to make their process competitive with the conventional process.

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Sarkar et al.7 and Bartels8 studied the effect of scale on the price of ammonia produced from

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biomass. Both predicted a steep rise in ammonia production price with a decreasing plant scale.

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Sarkar et al.7 predicted that an optimum biomass-to-ammonia plant would require 990,000 tpa of

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dry biomass.

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All the studies mentioned above indicate that ammonia production from biomass is feasible only

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at the production level of a world-scale ammonia plant. Smaller plants have to suffer the effects

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of economy of scale. However, biomass has a low energy density when compared to fossil fuels.

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As a rule of thumb, biomass should not be transported by trucks for more than 100 km, lest the

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cost of transportation would exceed the value of energy in the biomass.9 The biomass supply is

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also season-dependant and each biomass has a different harvesting time. This poses a significant

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challenge in obtaining a large scale supply of biomass throughout the year. The operation of

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large scale biomass plants is thus restricted to certain specific geographical locations where a

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large amount of biomass is economically available. In such a scenario, small scale biomass-

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based plants may be more feasible because they can be based on locally available biomasses.

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Additionally, they can cater to local ammonia demands effectively. A potential application

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could be found in the mining industry. Ammonia based mining explosives such as ANFO are

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conventionally produced on a small scale due to safety concerns and relatively small demands.

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The demand for such explosives originates from remote mining locations. A small-scale

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biomass-to-ammonia plant can thus easily meet the ammonia requirements for the production of

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explosives.

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This paper looks at the viability of small scale ammonia production from biomass. A scale of

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40,000 tpa of biomass intake is selected to buffer the sporadic biomass supply, while

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simultaneously accommodating the local ammonia demand. A conservative scale is chosen to

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assist the applicability of the process concept globally. The selected scale is much smaller when

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compared to conventional ammonia plants and the effect of the economies of scale is expected to

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make the plant less economically appealing. To overcome these effects, the design incorporates

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features from the ICI Leading Concept Ammonia® process. This small-scale process developed

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in the early 1990’s incorporated significant changes in terms of heat integration and process flow

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to suit small scale ammonia plants. The use of dual fluidized bed gasifier technology also

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contributes to a cleaner syngas. The details and contribution of these two design concepts are

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discussed in later sections. It may be noted that in earlier studies, when analysis was performed

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for small scale ammonia plants, these modifications were not considered and the process

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configuration was similar to the one used for large scale production.

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2. Concepts and methodology

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2.1 Process description

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Biomass gasification is a comparatively new technology with a small number of success stories

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at the industrial/commercial scale. Additionally, different gasifier configurations are available to

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suit the biomass feedstocks, output syngas quality and capacity range. With respect to the

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capacity range being considered in this study and the syngas yield needed to make a gasification

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process economically viable, fluidized bed biomass gasifiers are predicted to be the most

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economic.10 Among the fluidized bed gasifiers, dual fluidized bed gasifier configuration appears

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to be promising due to their potential to supply a clean syngas, in spite of the fact that it uses air


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as an oxidising agent instead to oxygen. One such gasifier is the 8 MW CHP gasifier at Gussing,

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Austria manufactured by Repotec.10 The gasifier has been successfully operating for heat and

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power applications since 2002. In a dual bed gasifier, the gasification and combustion take place

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in different vessels; the gasification is primarily carried out with steam, and combustion with air.

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The heat transfer takes place through the circulation of hot bed material between the two beds.

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The syngas produced is not significantly diluted with nitrogen, since steam is the sole gasifying

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agent. The methane content, however, is slightly higher. The heat produced from combustion of

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char is enough to sustain the gasification process and no external heat source is required.

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Furthermore, the flue gases exiting the combustor can be used for preheating of air, superheating

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of steam, drying of biomass and other heating applications. Other advantages of dual bed

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gasifiers when compared with conventional gasifiers according to Pfeifer et al.11 are:

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•

Low tars due to steam gasification

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•

No oxygen demand to obtain nitrogen free syngas

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Compact design and easy feeding of biomass

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•

A wide range of feedstock can be gasified

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•

The bed material can be used as a catalyst

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This study looks at small scale production of ammonia from biomass. Conventionally, ammonia

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is produced on a large scale with production capacities greater than 1000 tpd. Scaling down the

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conventional ammonia flowsheet would burden the economics. This study builds upon the

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Leading Concept Ammonia (LCA) ® process, which substantially contrasts with the

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conventional highly integrated large-scale ammonia plant.12 It is a small-capacity natural gas

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based ammonia plant that can compete with modern large-capacity natural gas based ammonia

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plants in terms of energy consumption and specific investment. The process energy consumption


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of the LCA process is 29.3 GJ/t NH3 for a natural gas feedstock, which compares well with 28

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GJ/t NH3 for state-of-the-art large scale technology.1,13 Furthermore, the technology competes

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well with large-scale productions in terms of economics. The design features responsible for

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improved economics are:13

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•

Milder Primary Reformer conditions

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•

A Gas-Heated Reformer

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•

An Isothermal CO Shift Reactor

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Pressure Swing Absorption (PSA) for CO2 removal

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•

A low pressure ammonia synthesis loop

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The PSA also helps in maintaining the H2:N2 ratio as well as removal of a portion of CO and

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CH4. This isothermal shift and PSA lead to robust plant operations and makes the process less

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sensitive to catalyst performance and plant upsets. The plant design is modular, eliminating any

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unwanted interactions and interdependencies. This gives the plant the flexibility to operate

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efficiently at reduced feed rates. Additionally, start up and shut down steps in small-scale plants

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are easier and faster, thus reducing operation costs.12

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2.2 Process simulation

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A steady state flowsheet for ammonia production from biomass was prepared in ASPEN Plus®

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simulation software (Figure 1). The software employs a sequential modular approach to compute

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the heat and mass balance. The ASPEN Plus® software has been utilized by several authors of

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which Gilbert et al.5, Tock et al.6 and Spath et al.14 are just the most recent for similar biomass

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processes. The package offers reliable thermodynamic data, realistic operating conditions and

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rigorous equipment models to simulate a variety of processes. In this work, we use the Redlich


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Kwong Soave Boston Mathias (RKS-BM) property method to model the entire flowsheet. This

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property package is used for processes that involve non-polar and real components, such as gas

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production, gas processing and hydrocarbon separation.15 The same property method is also

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recommended by ASPEN Plus® to simulate ammonia plants.

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Figure 1. Process flowsheet used in the present work for simulation and optimization

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Gibbs or equilibrium reactors were used to model the flowsheet reactors. These reactors work on

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the principle of minimizing the Gibbs free energy, thus, predicting the equilibrium composition.

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The choice of this reactor model is based on the fact that the undergoing reactions are known to

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reach close to equilibrium.16 Initially, a flowsheet replicating the small scale ICI LCA process

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was developed. The model was validated against the plant data reported in open literature.2,13,17

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An error of 2.1 % was realized in the final ammonia production. The flowsheet was then

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modified to incorporate the biomass gasification process and subsequent gas conditioning,

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replacing the natural gas feed used originally. The LCA process features that were included in


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the proposed biomass-to-ammonia process, are isothermal shift, PSA, low pressure ammonia

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loop and recycle of purge ammonia (Figure 1). The results were once again compared with

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different models available for ammonia and hydrogen production from biomass14,18 and were

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found to be in good agreement. Model details on the process are given in section 3.

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The heat and mass balance data obtained from the ASPEN Plus® flowsheets was used to predict

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the economics of the production of ammonia from biomass. An MS Excel® worksheet imported

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the data from the ASPEN Plus® flowsheet. The capital and operating costs were approximated,

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based on the correlations given by Turton et al.19 These correlations predicted the cost of major

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pieces of equipment based on mass flowrates, heat duties, operating pressures and the material of

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construction. Chemical Engineering Plant Cost Index (CEPCI)20 was used to account for the

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effects of inflation. This can be seen in Equation 1.

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, = ,

(1)

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where, CBM refers to the bare module cost, n is capacity exponent, Si are the scales of piece of

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equipment and Ii are the inflation indices for different years. Once the bare module costs were

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calculated, they were added together to calculate the grassroots-level costs by including

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contingency costs, other miscellaneous costs and the cost for auxiliary facilities, using

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Equation 2.

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= ∑ , + ∑ ,

(2)

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where, CGR is the total grass roots level costs, A1 is a factor used for contingency and fee costs,

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A2 is a factor used for auxiliary facilities cost, CBM is the bare module cost at process conditions

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and CBM0 is the bare module cost at ambient pressure and with carbon steel as fabrication

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material. The manufacturing costs were calculated for three main categories, namely, the direct

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cost including raw material, labor and maintenance cost; fixed costs that took care of

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depreciation, taxes and insurance; and general expenses that represented administrative costs,

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distribution and selling costs, and research and development. The economic performance of a

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plant is primarily based on the initial capital investment and the annual operation costs. The

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annualised capital cost along with the operating cost was used to calculate the total production

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cost, and hence, the minimum ammonia production cost.

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2.3 Flowsheet optimization

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In process design, the aim is to propose the best process possible in terms of efficiencies, in

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particular, positive economics and low environmental impacts. Sometimes, these objectives are

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conflicting or interfering, such that improvement in one objective leading to a decline in the

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other. Maximizing carbon sequestration while simultaneously minimizing production costs is a

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typical problem faced by both chemicals and energy industries.21 In the past, such systems were

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optimized by placing various weights on different objectives and optimising the combined single

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objective. Multi Objective optimisation (MOO) on the other hand, gives a range of optimal

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solutions called Pareto-optimal solutions for conflicting objectives, in which each solution is said

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to be non-dominated by other solutions. If a solution thus obtained is better for one objective, it

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is inferior for the other objective. No solution on a Pareto curve is thus better than any other

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point with respect to all the objectives. Thus, it presents a pool of equally good points, and the

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user can choose the optimal points best catering to his/her needs.


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The Non-Sorting Genetic Algorithm (NSGA) II algorithm developed by Deb et al.22 has been

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used effectively to optimize similar systems.23 The algorithm has been coded in Visual Basic

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(VB)/MS Excel® and linked to ASPEN Plus® by Rangaiah et al.24,25 The same code has been

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modified for this study, to optimize the economic parameters as well. The optimization

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methodology has been illustrated in Figure 2. The VB interface supplies design variables to the

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ASPEN Plus flowsheet and based on the flowsheet solutions, the process economics are

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calculated. The VB interface then extracts values of objective functions to be optimized. Based

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on the extracted values, new generations of solution are created. The generations, so created in

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the cyclic process, undergo crossover and mutation for steady improvement of solutions. The

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ASPEN Plus flowsheets that do not converge due to infeasible input are neglected by the code.

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The production of better generations continues till a termination criterion is met. The termination

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criteria used in this study is the maximum number of generations.


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Figure 2. MOO methodology used for the analysis of the two different flowsheets of interest

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3. Model Description

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3.1 Process Flowsheet

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This section describes the ASPEN Plus® model and associated economic calculations for the

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biomass-to-ammonia process. The simulation starting point was chosen to be the biomass

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feeding to the dual bed gasifier. The flowsheet was simulated for eucalyptus as the feedstock

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biomass. The ultimate and proximate analyses of the eucalyptus biomass are presented in Table

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1. A compartment based semi-detailed kinetic model proposed by Arora et al.26 was used to

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model the dual fluidised bed gasifier (Figure 3). The gasifier model has been validated against

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the plant data for the CHP gasifier.26 The gasifier and the combustion chambers were modelled

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as separate reactors and the heat integration between the two was made possible by circulation of

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bed material. The output syngas has a high hydrogen content resulting from steam, which is used


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as a gasifying agent. The same also results in higher methane content, which would needs to be

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reformed.

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Table 1. Proximate and ultimate analyses of biomass feedstock

Proximate analysis (weight %) Moisture (as received) Fixed Carbon (dry basis) Volatile Matter (dry basis) Ash (dry basis) Ultimate Analysis (weight %) (dry basis) Carbon Hydrogen Nitrogen Chlorine Sulfur Oxygen Ash

Biomass 27 16.20 83.80 1.16 48.00 6.20 0.20 0.05 0.02 45.60 1.16

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Figure 3. Compartment model for the FICFB gasifier

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The hot syngas needs to be reformed for the removal of methane as well as higher hydrocarbons

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including tars. This study looks at two different approaches for the reforming reactions

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(Figure 4). The first approach utilizes an Autothermal Reformer (ATR), where both steam and


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air are used as oxidizing agents. The heat for the endothermic reforming reactions is provided by

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the combustion reactions that occur in the presence of air and no external heat source is required.

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Similar to a secondary reformer used in conventional ammonia plants, the heat and nitrogen

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requirements are met by adjusting the air supply. In the second approach, steam alone is used as

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the oxidizing agent and reforming occurs using conventional Steam Methane Reforming reactor

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(SMR). The heat required to sustain the endothermic reactions is supplied through the

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combustion of an additional fuel, namely, natural gas.

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Figure 4. Representative flowsheets for ATR and SMR configurations

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The reforming of syngas is followed by gas cleaning using filters, scrubbers and coolers for

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removal of particulates and inorganic impurities such as halides, alkali and sulfur compounds.

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The study assumes the use of hot gas cleaning technology for syngas cleaning, due to better

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efficiency.27 The technology, however, is in the development stage. The gas is then compressed


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using a three-stage compressor. The compressed gas then enters an isothermal shift reactor for

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conversion of CO to CO2 by the water-gas shift reaction. The study uses a single shift reactor as

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opposed to two shift reactors used in conventional plants. This can be justified, since the absolute

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carbon-monoxide slip would be small due to the smaller flow-rates and can be taken care of by

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the PSA and methanator downstream.16 Thus, an entire reactor vessel in the form of an additional

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shift reactor can be avoided. Furthermore, the heat produced in the shift reactor is used to

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generate steam that is used as reactant in the process for reforming and in the shift reactor.

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The desaturator follows the shift reactors and reduces the moisture content of the syngas. The gas

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then enters a PSA to remove CO2 and leftover CO, CH4 and inerts. The PSA can also be used to

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remove excess nitrogen that results from the excess air used in the ATR. Thus, the PSA plays a

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critical role in maintaining the appropriate H2-to-N2 ratio. The hydrogen recovery of the system

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was taken as 87 %. PSA is the only place where inerts are removed from the process. The PSA

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flue-gas also acts as a low calorific value fuel. After this, the cleansed syngas enters a methanator

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for removal of oxides, namely CO2 and CO, which are poisonous to ammonia synthesis catalysts.

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They are converted into methane by the Methanation reactor. After this, a molecular sieve drier

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is used to strip the gas of any remaining moisture content. In the SMR approach, an ASU is used

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at this point to supply the required nitrogen. The syngas now enters the Haber’s Bosch synthesis

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loop. Based on the ICI LCA process, the ammonia synthesis loop is operated at low pressures.

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This helps in achieving compressor efficiencies similar to the large-scale ammonia plants.12 The

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process uses a tube-cooled counter-current reactor that contains a low pressure ammonia

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synthesis catalyst and operates at temperatures upto 400 0C.17 To prevent the accumulation of

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methane, a purge is introduced in the synloop. The purge is recycled and mixed with the syngas


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that enters the reformer. The complete flowsheet for the process (ATR configuration) can be

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seen in Figure 1. The representative flowsheets for ATR and SMR process configurations are

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shown in Figure 4. The operating parameters for major unit operations are shown in Table 2.

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These parameters are based on the ICI process and are found to have delivered better results in

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small-scale operations. Stream results for the two flowsheets are reported in Table 3 and Table 4.

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Table 2. Operating parameters for unit operations

Parameter

Nominal

Range

Steam-to-biomass ratio

0.4

0.3-1

Gasifier temperature (°C)

792

750-850

Combustor temperature (°C)

894

850-950

Shift temperature (°C)

265

220-400

PSA H2 recovery (%)

87

80-95

Synthesis reactor pressure(Bar)

82

80-300

Synthesis temperature (°C)

380

300-500

H2/N2 ratio

3

2-4

Synthesis loop purge (%)

2.6

2-10

316 317

Table 3. Stream table for biomass-to-ammonia process (ATR configuration) ATR case Stream No.

1

2

3

4

5

6

7

8

9

Temperature (0C)

792.3

807

250

265

35

35

45

-30

30

Pressure (bar)

1

1

30

29

28

28

27

75

75

Mass Flow (kg/h)

6642.08

10007.7

8478.48

14478.5

10245

3070.36

3025.54

2659.71

365.57

CH4

0.079

289 PPM

342 PPM

198 PPM

282 PPM

223 PPM

0.004

0.002

0.035

H2O

0.288

0.164

0.011

0.297

0.002

0.003

-

-

-

H2

0.278

0.366

0.433

0.380

0.540

0.743

0.742

0.001

0.669

Mole Fraction


16

Page 17 of 39

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


CO

0.142

0.193

0.228

0.004

0.005

0.004

-

-

-

CO2

0.161

0.092

0.109

0.192

0.272

-

-

-

-

N2

0.026

0.182

0.216

0.125

0.178

0.248

0.252

527 PPM

0.260

NH3

-

34 PPM

-

-

-

-

-

0.995

0.022

318 319 320 321 322 323 324

Table 4. Stream table for biomass-to-ammonia process (SMR configuration) SMR case Stream No. 0

Temperature ( C)

1

2

3

4

5

6

7

8

9

792.3

850

250

265

35

35

45

-30

30

Pressure (bar)

1

1

40

39

38

Mass Flow (kg/h)

6642.08

7008.33

5912.87

11912.9

8019.56

889.39

38

37

75

75

842.328

3066.32

366.03

CH4

0.079

438 PPM

503 PPM

277 PPM

391 PPM

326 PPM

0.006

0.002

0.040

H2O

0.288

0.137

0.009

0.292

0.001

0.002

-

-

-

H2

0.278

0.491

0.564

0.473

0.667

0.968

0.970

148 PPM

0.704

Mole Fraction

CO

0.142

0.264

0.303

0.005

0.007

0.006

-

-

-

CO2

0.161

0.068

0.078

0.205

0.289

0.000

-

-

-

N2

0.026

0.040

0.045

0.025

0.035

0.024

0.024

776 PPM

0.234

NH3

-

19 PPM

-

-

-

-

-

0.997

0.022

325 326

3.2 Economic Parameters

327

In comparing different processes for manufacturing NH3, the profitability of the operation is

328

bound to be one of the main deciding factors. As stated earlier, this study looks at two different

329

configurations for small-scale ammonia production from biomass and predicts the ammonia gate

330

price for both the configurations. One of the main objectives in optimizing the process is to

331

derive maximum economic benefit. Eastern Australia was chosen as the plant location with the

332

economic parameters that are appropriate for this location. The eucalyptus plantation needs an

333

establishment period of 4 - 5 years.28 After this period it can be grown as a short rotation coppice


17


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Page 18 of 39

334

with harvesting taking place every 2 - 4 years; the yield is approximately 5 - 10 dry

335

tonne/hectare/year.28 It is reported that the delivered cost of energy crops is made up of 25

336

percent production cost of crops, 50 percent harvesting and 25 percent transport. The cost of

337

biomass is highly variable and the cost per tonne for eucalyptus has been reported to be from $70

338

to $150 per oven dry ton (odt).28 The present study takes the biomass gate price in the middle of

339

this range at $ 100/odt.

340 341

Initial capital investment and annual operating costs are the two main parameters that determine

342

the economic viability of a manufacturing process. These parameters were calculated based on

343

the methodology suggested in section 2. Gas Hourly Space Velocity (GHSV) was used as a basis

344

for sizing of major reactor vessels. The GHSV data reported in open literatures29,30,31 were used

345

for sizing the reformer, the shift reactor, the methanator and the synthesis reactor. The basic

346

reformer cost, air separation cost and the gasifier system cost were taken directly from

347

literature.10,14,32 The heat exchangers were modelled as floating head shell and tube heat

348

exchangers. A methodology proposed by Riberio et al.33 was used to size the PSA beds, which

349

was assumed to have six beds. Adsorption isotherms reported by Jee et al.34 were employed for

350

the above methodology. The ammonia loop was sized on the basis of approach suggested by

351

Morgan.32 Stainless steel was considered to be the material of construction for the entire plant

352

since the presence of hydrogen in syngas can cause serious damage to carbon steel

353

equipment.32,35 A material factor for stainless steel was used to reflect the increase in the material

354

costs. Similarly, pressure factors were used to predict the equipment costs that are operated at

355

high pressures.

356


18

Page 19 of 39

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


357

In order to calculate the operating costs, the costs of the biomass feedstock, electricity, different

358

catalysts, adsorbent, bed material, ash disposal, and boiling feed water were considered. The

359

electricity consumption by ASU was approximated based on the correlations suggested by Hu et

360

al.36 The labor cost was taken to be $ 90,000 per person per annum.37 The direct supervisory

361

labor was assumed to be 18 % of the operating labor cost. Maintenance and repairs were

362

considered to be 4 % of the Fixed Cost of Investment (FCI) or Grassroots level costs. Local taxes

363

were assumed to be 1.4 % of the FCI.19 The operating supplies, laboratory charges, patents and

364

royalties, plant overhead costs and general manufacturing costs were estimated by correlations

365

suggested by Turton.19 All the costs are reported in terms of the 2014 US Dollar. The major

366

assumptions in calculating the ammonia production cost are listed in Table 5.

367

Table 5. Economic parameters.

Parameter

Value

Biomass input(odt/day)

120

Project life (years)

20

Plant availability (%)

90

Discount rate (%)

15

Biomass price($/GJ)

5.6

Natural gas price ($/GJ)

5.5

Electricity price ($/GJ)

41.7

368 369

4. Flowsheet Optimization

370

The two configurations of interest (ATR and SMR) differ in the manner in which reforming of

371

hydrocarbons present in syngas is carried out. The first configuration has been named ATR since

372

it employs an autothermal reactor, and the energy requirements for the endothermic methane

373

reforming are provided by burning a part of syngas. The other configuration has been named


19


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

Page 20 of 39

374

SMR since it uses the conventional steam methane reformer. Here, an external heat source such

375

as natural gas combustion is used to provide the necessary heat. The nitrogen needed is supplied

376

as a part of the air in the flowsheet that incorporates ATR. In the SMR flowsheet, the

377

requirement of nitrogen is provided by the use of an ASU. The use of enriched air from ASU for

378

the gasification process as well as the use of gasification syngas as a fuel for production of

379

external heat for the SMR reactor has been left as a future exercise.

380 381

In order to predict which of the two configurations was better, a fair comparison between both

382

the flowsheets was required, which in turn requires each of the flowsheets to be optimized. The

383

capital and operating costs of both the flowsheets were minimized using the MOO methodology.

384

ATR/SMR temperature, air supply (only for the ATR case) and gas conditioning pressure The

385

MOO was performed for 30 generations and a population size of 100. The gas conditioning

386

pressure was found to have a major impact on the plant economics. The influence of other

387

variables was not found to be significant. An increase in gas conditioning pressure for both

388

configurations led to an increase in the capital cost. The operating cost, however, decreased. The

389

results are illustrated in Figure 5. A higher pressure led to an increase in the cost of equipment,

390

and hence, the final capital cost. On the other hand, an increase in pressure also reduced the

391

combined load on compressors. This further led to lower electricity consumption, and hence, the

392

operating cost. The change in the final ammonia manufacturing cost was not significantly

393

different for the different gas conditioning pressures (Figure 6). Nevertheless, the results show

394

minima in ammonia cost, observed at 30 Bar for the ATR system and that at 40 Bar for the SMR

395

system. These pressures were used to simulate the respective configurations.


20

750 740 730 ATR

720

SMR

710 700 400

397

420

440

460

480

500

Annualised Capital Cost ($/t NH3)

396

Figure 5. Comparison of optimized capital and operating costs (Pareto curves)

1170 Ammonia price($/t NH3)

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


Operating Cost ($/t NH3)

Page 21 of 39

1165 1160 1155

ATR

1150

SMR

1145 1140 20

30 40 50 Gas Conditioning Pressure (Bar)

398 399

Figure 6. Comparison of the effects of gas conditioning pressure on the ammonia price for ATR and SMR

400 401

The steam-to-carbon ratio used for the shift reactor was also found to influence the minimum

402

ammonia production cost. An increase in steam leads to a higher hydrogen concentration after

403

the shift reactor. This further increases the final ammonia production for the same input, thus,

404

bringing down the ammonia production cost. This trend however stabilized after reaching a

405

steam-to-carbon ratio of 2; a further increase in steam had only a very small effect on the


21


406

ammonia production cost. The results are illustrated in Figure 7. It may also be noted that a

407

carbon deposition on shift catalysts is experienced at low steam-to-carbon ratio. The deposited

408

carbon, reacts with catalyst to form iron carbide, which is a potential Fisher–Tropsch catalyst and

409

may lead to the formation of methane and other hydrocarbons.31

Ammonia price($/t NH3)

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

Page 22 of 39

1240 1220 1200 ATR

1180

SMR 1160 1140 0

1

2

3

Steam to Carbon ratio for Shift Reactor 410 411

Figure 7. Comparison of the effects of steam input to Shift Reactor in ATR and SMR

412

5. Energy Integration

413

The two flowsheet configurations were based on the LCA design philosophy and so was the heat

414

exchange network. The ICI LCA process for ammonia production from natural gas uses a heat

415

exchange reformer, wherein, heat from the cooling of secondary reformer syngas is utilized to

416

heat the primary reformer. However, the absence of primary reforming in the biomass-to-

417

ammonia flowsheet created a scope for additional heat integration. Sensible heat could be

418

potentially extracted from the cooling of the syngas outlet from SMR/ATR. The LCA flowsheet

419

utilized this heat for primary reforming or gas-heated reforming. Additional heat could also be

420

extracted from the ammonia synthesis loop (also available in the LCA plant) and the cooling of

421

the combustor flue gas. The PSA off-gas can also be used as a low calorific value fuel. One


22

Page 23 of 39

422

potential application is the preheating of air for the ATR configuration. A Grand Composite

423

Curve (GCC), that reflects the heat quantities available at different temperature intervals,38 is

424

plotted for the two flowsheets in Figure 8. It shows that both the systems have very similar GCC

425

profiles. This extra heat can be effectively used for the production of electricity. 1000.00 800.00 ATR

Temperature 0C

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


600.00

SMR

400.00

Steam Composite Curve

200.00 0.00 0.00 426

5.00

10.00 Heat MW

15.00

20.00

427

Figure 8. Grand Composite Curves for ATR and SMR and Steam Composite Curve

428

The GCC of both the plant configurations reflect the scope for utility steam and electricity

429

generation. Electricity, if taken from grid contributes to almost one-fifth of the production cost of

430

ammonia. Electricity is utilized mainly to run the compressors. Thus, the ammonia production

431

cost could be reduced if the waste heat available is utilized for electricity production. The

432

flowsheets of both process configurations were altered to incorporate electricity production from

433

waste heat. The heat available in the plant was used to produce steam at 40 bar and superheated

434

till 450 0C. A back pressure turbine was used to generate electricity from this steam. A part of

435

the let-down steam was used in the gasifier and the drying of the biomass. A condensing turbine

436

was employed to produce additional electricity from the remaining steam. The process diagram

437

for electricity production is shown in Figure 1. The ATR and the SMR processes generated 3.6


23


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

Page 24 of 39

438

MW and 3.7 MW of electricity, respectively. This brought down the final electricity requirement

439

from grid for the ATR and the SMR process to 0.9 MW and 1.1 MW of electricity, respectively.

440

The steam composite curve is shown in Figure 8.

441

A better process optimization could be performed if heat integration is also included in the MOO

442

framework.39 The formulation of GCC and predicting the optimum steam generation for every

443

subsequent ASPEN Plus flowsheet adds significantly to the complexity of the optimization

444

framework and it is recommended that this be undertaken in the future analysis. However, it will

445

be shown in the results section that electricity production makes only a small contribution to the

446

final ammonia production cost.

447 448

6. Results and discussion

449

The economic analysis for the optimized flowsheets of both plant configurations is now

450

presented. The initial capital investment is divided into five different plant sections, namely,

451

gasifier cost (including feedstock treatment), gas conditioning cost, ammonia synthesis loop cost,

452

air separation unit cost and steam turbine (and generator) cost. The annualised capital cost break-

453

down is illustrated in Figure 9. The gasifier cost was found to be the biggest contributor to the

454

total capital investment. The annualised gasifier capital cost is larger for the ATR system

455

because the ammonia output is lower in the ATR case, and the costs reported are normalised for

456

ammonia production. The next major cost was gas conditioning for both the flowsheet

457

configurations. This cost is slightly higher in the case of SMR because of higher pressure

458

operation. A similar variation is also observed in the cost of ammonia synthesis loop. By

459

contrast, the ASU cost is only applicable to the SMR flowsheet.


24

Page 25 of 39

600.00

500.00 Annualised Capital Cost ($/t NH3)

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


400.00 Steam turbine ASU 300.00

Ammonia Synthesis Gasifier

200.00

Gas Conditioning

100.00

0.00

460

ATR

SMR

461

Figure 9. Capital cost break-down for biomass-to-ammonia plant for ATR and SMR

462

To understand the cost contribution of individual items, the costs of some of the major pieces of

463

equipment are reported in Table 6. SMR/ATR reactors, PSA, compressors and heat exchangers

464

are the major cost contributors. The PSA system is a major cost for both the process

465

configurations. An additional compressor needed for the ASU makes the compressors cost

466

higher for the SMR system. The complex tubular configuration of the SMR reactor makes its

467

cost higher than that of the ATR reactor. The flow of syngas is greater for the ATR configuration

468

since the nitrogen required for ammonia synthesis is introduced in the ATR. However, the effect

469

of a high flowrate on the size and cost of pieces equipment is not very significant.

470 471 472


25


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

473

Page 26 of 39

Table 6. Bare module cost of select operations.

Unit Operation

Cost(Million $) ATR

SMR

Compressors

4.82

4.97

SMR/ATR

0.86

1.37

Shift

0.22

0.29

Desaturator

0.06

0.06

PSA

4.30

5.25

Methanator

0.05

0.06

Heat Exchanger

2.37

2.67

Synthesis Reactor

0.71

0.82

ASU

-

8.68

474 475

The operating costs for the plant were found to be dominated by the biomass feedstock costs

476

followed by the maintenance and labour costs. Figure 10 presents the cost build-up for the final

477

ammonia production cost. Capital cost recovery is predicted to be the largest contributors to the

478

total cost of manufacturing. The higher ammonia production in the SMR configuration justified

479

its higher capital cost. This also explains the reason for the lesser contribution of the biomass

480

feedstock cost and electricity cost to the total ammonia production cost, when compared to the

481

ATR system. Thus, both the plant configurations gave almost similar ammonia manufacturing

482

costs. The ATR system, however, has a slight cost advantage to offer.


26

Page 27 of 39

1400.00 Annual Investment

Ammonia total cost build-up ($/t NH3)

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


1200.00

Plant Overhead

1000.00

Taxes and Insurance Maintenance

800.00

General Manufacturing Expenses

600.00

Labour cost

400.00

Catalysts and Utilities

200.00

Electricity Biomass Feedstock

0.00 483

ATR

SMR

484

Figure 10. Ammonia cost build-up for ATR and SMR

485

The capital costs for the ATR and the SMR configurations were calculated to be $ 62.9 M and

486

$ 74.4 M, respectively. The ammonia production for the ATR configuration was 65 tpd and for

487

the SMR configuration, it was 74 tpd. The biomass input was constant at 120 tpd for both the

488

process configurations. The ammonia production costs were calculated to be $ 1153/t NH3 and $

489

1172/t NH3 for the ATR and SMR systems, respectively. A sensitivity analysis for the ammonia

490

costs was carried out and it was found that the biomass price and discount rate had a major

491

impact on the ammonia costs. The sensitivity analysis results are presented in Figure 11 and

492

Figure 12. Additionally, different studies predicting ammonia production costs from biomass

493

feedstock, show significant variation while choosing these two parameters. Thus, the sensitivity

494

analysis helped to perform a fair comparison with different economic analyses. The results


27


495

predict that 1 % variation in discount rate alters the ammonia production cost by approximately $

496

25/t NH3 for both the process configurations. Similarly, a variation of $ 1 in biomass price led to

497

a variation of approximately $ 2/t NH3 for the ATR system and approximately $ 1.8/t NH3 for

498

the SMR system. These different sensitivities result from a higher throughput of the SMR

499

system. The ATR configuration appears to have a slight economic advantage over the SMR

500

configuration in most of the cases. However, for a lower discount rate and a higher biomass

501

price, the SMR configuration offers a small economic advantage because its cost is less sensitive

502

to the biomass price.

1500 Ammonia price ($/t NH3)

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

Page 28 of 39

1400 1300 1200 1100 1000 900

Biomass price $ 50/odt

800


700


600 0

5

10 15 Discount rate (%)

20

25

503 504

Figure 11. Sensitivity analysis for biomass feedstock price for the ATR configuration


28

Page 29 of 39

1500 Ammonia price ($/t NH3)

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


1400 1300 1200 1100 1000 900 Biomass price $ 50/odt Biomass price $ 100/odt

800 700

biomass price $ 150/odt

600 0

5

10 15 Discount rate (%)

20

25

505 506

Figure 12. Sensitivity analysis for the biomass feedstock price for the SMR configuration

507

The results above were compared with other studies that undertake the economic analysis of

508

ammonia production from biomass (Table 7). The comparison highlighted the economic

509

advantages resulting from the process concepts of the LCA plant. In the study carried out by

510

Gilbert et al.5 for their 1200 tpd ammonia production plant, they have reported an ammonia

511

production cost of $ 386/t NH3. However, the biomass feedstock price assumed is approximately

512

$ 60/odt, which is on the lower side. Tock et al.6 for their 1187 tpd ammonia production plant,

513

have predicted an ammonia production cost of $ 968/t NH3. This cost also takes into account the

514

carbon capture cost. The discount rate assumed by these authors is 6 %. Another similar study

515

carried out by Anderson et al.40 reported a cost of approximately $ 750-1200/t NH3 for a 700 tpd

516

ammonia plant. Bartels8 had predicted an ammonia production cost of $ 488/t NH3 for a 1022

517

tpd ammonia plant and $1519/ t NH3 for a 10 tpd plant. The ammonia synthesis loop capital

518

predicted by this study is approximately one third of their predicted value at a similar scale.

519

These results indicate that there is reasonable variation in the predicted estimates. The economic


29


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

Page 30 of 39

520

estimates are very sensitive to assumptions such as the discount rate and plant life. Additionally,

521

operating costs such as biomass feedstock costs are dictated by the geographical location under

522

study. The effect of these variations has been depicted in the sensitivity analysis results presented

523

in Figure 11 and Figure 12.

524

Table 7. Ammonia production cost comparison

Study

Ammonia Biomass Production Price Scale (tpd) ($/odt)

Discount Ammonia Rate (%) Price($/t NH3)

Remarks

Bartels8 (2008)

10

16.5

8

1519 Model details not available

Bartels8 (2008)

1022

46.3

8

488 Model details not available

Gilbert et al.5 (2013)

1200

60

20

500 Low feedstock cost

Tock et al.6 (2014)

1187

260

6

968 Low discount rate

700

70-280

10-20

750-1150

-

65

100

15

1153

-

Anderson et al.39 (2014) This Study

525 526

Most of the economic predictions reported above have been carried out at world-scale ammonia

527

production levels, since, reducing the scale would further increase the production cost of

528

ammonia. A study by Sarkar et al.7 reports that ammonia production costs rise exponentially at

529

scales below 1000 tpd of biomass use. In our study, however, the cost advantages of a world-

530

scale ammonia plant can be achieved at plants of much smaller scale too. The comparison is

531

illustrated in Figure 13. Another rationale behind selecting a conservative scale of 65 tpd for the

532

biomass-to-ammonia plant was that, any small scale ammonia production facility would

533

generally be larger than this scale and benefit from the economies of scale. Furthermore, if the

534

designed flowsheets were used to produce ammonia at a scale of 1220 tpd, the ammonia

535

production cost is approximated to be $ 743/t NH3 for the ATR configuration and $ 742/t NH3


30

Page 31 of 39

536

for the SMR configuration. The capital costs were calculated to be $ 585 M and $ 688 M for the

537

ATR and SMR sheets, respectively. The difference in capital costs is mainly attributed to the

538

ASU. These costs, however, can be further reduced by use of technologies more suited to a

539

higher scale, such as use of entrained flow gasifiers. It can thus be concluded that the

540

incorporation of small-scale features of the LCA process and the use of dual fluidized bed

541

gasifiers for the biomass-to-ammonia process, have helped to reduce the effect of economies of

542

scale. This innovative combination of new and old technologies at front end and tail end of plant

543

respectively, can help in worldwide application of biomass-to-ammonia concept, which would

544

otherwise be constricted with biomass feedstock supply.

4000 Ammonia Price ($/t NH3)

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


3500 3000 Sarkar et al.

2500

Bartels

2000

Gilbert et al.

1500

Tock et al.

1000

Anderson et al. Present work

500 0 10

100 1000 Ammonia Production Scale (tpd)

545 546

Figure 13. Comparison of the production prices of ammonia in various studies


31


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

Page 32 of 39

547

7. Environmental impacts

548

In the use of renewable sources of energy, the environmental impact of a plant plays a critical

549

role. Ammonia plants are often criticized for their huge carbon footprints. The use of biomass as

550

a feedstock can help solve this problem, since most of the CO2 emitted by a biomass-based plant

551

is biogenic. This means that the CO2 released is later consumed by the growing biomass

552

feedstocks and no extra CO2 is added to the atmosphere. The biogenic emissions from the

553

gasifier and the PSA have been neglected for calculating the total CO2 emissions. The Global

554

Warming Potential (GWP) of the two flowsheets was calculated based on the Recipi

555

methodology, with a hierarchist perspective of 100 year timeframe. The GWP was estimated

556

based on the ASPEN Plus® simulation data and Ecoinvent® databases, normalised for

557

Australian conditions. SimaPro® life cycle assessment software was used to calculate the cradle-

558

to-grave life cycle assessment for the biomass-to-ammonia process for the two flowsheets. The

559

system boundary included biomass production, transportation, gasification leading till ammonia

560

production and utilities including electricity, water, bed material and catalysts (Figure 14). 1 kg

561

of NH3 produced was chosen as the functional unit. The total CO2 emissions adding to Global

562

Warming Potential (GWP) for the ATR and SMR systems were estimated to be 0.588 kg CO2

563

eq./kg NH3 and 0.841 kg CO2 eq./kg NH3, respectively. The CO2 emissions results are shown in

564

Figure 15. These emissions are much lower than the emissions from a typical natural-gas fed

565

ammonia plant. A conventional natural gas based ammonia plant in Australia has GWP of 1.84

566

kg CO2 eq./kg NH3 for similar assumptions. The GWP for a partial oxidation based ammonia

567

plant has a GWP of 3.09 kg CO2 eq./kg NH3 .The fossil fuel-based electricity production seems

568

to be the major CO2 emitter for the proposed flowsheets.


32

Page 33 of 39

569 570

Figure 14. System boundary for GWP impact assessment

0.9 0.8

CO2 emmissions Kg CO2/Kg NH3

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


0.7 0.6 Others FICFB Gasification

0.5

Eucalyptus cultivation

0.4

Transportation Fuel(Natural Gas)

0.3

Electricity Production

0.2 0.1 0 571 572

ATR

SMR

Figure 15. CO2 emissions for the biomass-to-ammonia process for ATR and SMR configurations


33


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

Page 34 of 39

573 574

8. Conclusions

575

A techno-enviro-economic analysis of small-scale ammonia production using biomass as

576

feedstock has been demonstrated. The proposed route leverages the inherent benefits of small-

577

scale ammonia production and the use of dual fluidized bed biomass gasification technologies.

578

The predicted economics has been found to be comparable to large-scale biomass-to-ammonia

579

production plants. This novel amalgamation of relatively new gasification method particularly

580

suited to biomass at small scale as the front end of the plant, with a largely over-looked LCA

581

process for the tail-end of the plant, can help in practical realization of the environment friendly

582

low carbon footprint biomass-to-ammonia process. Two different flowsheets namely, ATR and

583

SMR configurations have also been compared. Although a similar ammonia manufacturing cost

584

is predicted for both flowsheets, the capital and operational cost of both configurations differ

585

significantly. These small-scale ammonia production flowsheets can effectively meet the

586

ammonia requirements for small countries or isolated applications such as mining activities.

587

Being small scale, the required biomass feedstock could also be sustainably produced. The life

588

cycle assessment predicts that using biomass feedstock leads to GWP reductions of 54-68 %,

589

when compared to conventional ammonia plants. The emissions can further be reduced by

590

employing carbon sequestration. However, this could make the economics less attractive. The

591

trade-off between ammonia economics and greenhouse gas emissions requires further

592

investigation.

593 594

Acknowledgements


34

Page 35 of 39

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


595

The authors would like to thank Orica Ltd. for funding the project through IITB-Monash

596

Research Academy.

597 598

Supporting information

599

The ASPEN Plus® simulation files (ATR.bkp and SMR.bkp) for the ATR and SMR variants of

600

biomass-to-ammonia process have been uploaded for reference.

601 602

References

603 604

(1)

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(4)

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692 693

Nomenclature

694

$/t - U.S. dollars per ton


37


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

695

$/GJ - U.S. dollars per gigajoule

696

AER - Absorption enhanced reforming

697

ANFO - Ammonium nitrate/fuel oil

698

ASU - Air separation unit

699

ATR - Autothermal reformer

700

CBM - Bare module cost

701

CBMO- Bare module cost at base-case conditions

702

CGR - Grassroots level cost

703

CEPCI - Chemical Engineering Plant Cost Index

704

CO - Carbon monoxide

705

CO2 - Carbon dioxide

706

CH4 - Methane

707

C2H4 - Ethene

708

DFBG - Dual fluidized Bed Gasifier

709

FCI - Fixed cost of investment

710

GCC - Grand Composite Curve

711

GHG - Greenhouse gas

712

GHSV - Gas Hourly Space Velocity

713

GJ/t - Gigajoule per ton

714

GWP - Global warming potential

715

H2 - Hydrogen

716

ICI - Imperial Chemical Industries

717

km - kilometre

Page 38 of 39


38

Page 39 of 39

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


718

LCA - Leading Concept Ammonia

719

MOO - Multi Objective Optimization

720

M - Million

721

MW - Megawatt

722

N2 - Nitrogen

723

NH3 - Ammonia

724

NSGA - Non-Sorting Genetic Algorithm

725

O2 - Oxygen

726

odt - Oven dry tonne

727

PSA - Pressure Swing Adsorption

728

RKS-BM - Redlich Kwong Soave Boston Mathias

729

SMR - Steam methane reformer

730

tpd - Tonne per day

731

tpa - Tonne per annum

732

VB - Visual basic

733 734 735 736 737 738 739 740


39

Small-Scale Ammonia Production from Biomass: A ... - ACS Publications

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