advanced ammonia optimize

7/30/2019 Advanced Ammonia Optimize

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The Application of Model Based Predictive Controlfor Ammonia Plants

by

Robert Lin, Reinder de Boer, Bill Poe

GE Continental Controls, Inc.

Introduction

There are about 400 ammonia plants globally. Most of these plants utilizenatural gas as feedstock and fuel in the production of ammonia. Because of the

high degree of interaction between each processing unit, weather changes,catalyst deactivation, equipment degradation and fluctuations in feed gas flowand composition, ammonia plants are not always operated at their optimum

state. An advanced process control system is necessary to continuouslyoptimize the operation of the ammonia plant. This article discusses theapplication of advanced process control for ammonia plants and presents actual

data showing operational improvements.

The Ammonia Production Process and Plant Operation

Although several processes have been developed for ammonia production, thepredominant processing unit layout includes natural gas desulfurization, primary

reforming, secondary reforming, water-gas shift conversion, carbon dioxide

removal, methanation, synthesis gas compression, ammonia conversion and itsassociated separation loop, a refrigeration system and the steam system. A

simplified process flow diagram is shown in Figure 1.

Desulfurization

Natural gas delivered to the battery limits contains sulfur compounds which arepoisonous to the nickel catalyst in the primary reformer. The sulfur must be

removed from natural gas before it is introduced to the primary reformer.Depending on the sulfur content in the natural gas, one or more desulfurizer

catalysts are used. Normally, Zinc-Oxide catalyst performs as a final stage tokeep the remaining sulfur in the natural gas to a non-detectable level in theoutlet gas.

The operation of the desulfurizer mainly depends on the Zinc-Oxide catalystoperating temperature and the sulfur absorption capabilities of the catalyst.


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Primary Reformer

The primary reformer can be separated into two distinct sections:

a convection section, mainly used for heat recovery from the furnace

flue gas while preheating the reformer feed gas, and a radiant section where the actual reforming action takes place under

the influence of heat generated by dozens of fuel gas burners.

An auxiliary boiler is attached to the convection section to generate additionalhigh pressure steam required to drive the syngas compressor.

A typical design of the radiant section consists of a firebox with reformer catalysttubes arranged in parallel, single width rows, fired from both sides by burners

located in the furnace arch. The desulfurized natural gas first mixes with processsteam at a desired steam to carbon ratio and is then heated in the convection

section of the primary reformer. The heated mixed gas enters the tubes throughmanifolds above the furnace arch, flows downwards through the tubes, and iscollected in the insulated outlet manifold located near the floor. An outletmanifold is provided for each row of reformer tubes.

Combustion flue gases flowing downward in parallel to the catalyst tubes arecollected in flue gas tunnels at the bottom of the reformer. The flue gases then

pass through these tunnels into the convection section for heat recovery.

The operation of the primary reformer is critical. Not only is the primary reformer

the most expensive piece of equipment in the ammonia plant (about 30% of total

capital cost), but its operation will affect the whole operation of the plant.Unstable operation of the primary reformer will affect the steam system and the

performance of the downstream process units: secondary reformer, shiftconverters, methanator and synthesis loop. The primary reformer will greatlyaffect the ammonia production rate and overall energy consumption.

Several major factors affect the operation of the primary reformer, including:

feed gas flow and composition,

fuel gas flow and composition,

steam flow to the primary reformer, purge gas flow to the primary reformer,

weather changes (wind direction, rain storms, etc.),

catalyst deactivation,

mixed feed inlet temperature to the primary reformer, and

feed gas inlet pressure.


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The main objective of the primary reformer is to maximize methane conversionwith minimum fuel usage while preventing carbon deposition in the reformingtubes.

The operating objective of the primary reformer is to control the steam to carbon

ratio and the primary reformer exit temperature. Increasing the primary reformerexit temperature will increase the methane conversion but the fuel gasconsumption will increase as well and carbon deposition in the reforming tubesbecomes more likely. There is a trade-off between the increasing exit

temperature and increasing fuel gas flow. The exit temperature must becontrolled at a certain level. Due to the earlier mentioned factors anddisturbances, it is quite common to experience difficulties in maintaining the

primary reformer exit temperature at its desired value. Deviations of as much as8F from the target control point are regular.

The steam to carbon ratio is controlled excessively above the theoretical value

(in the region of 1.7) to reduce the carbon deposits in the reforming tubes.Lowering the steam to carbon ratio will decrease the energy consumption

greatly, but requires tight control of the reforming tube temperatures. Becauseof inconsistent secondary air openings between fuel gas burners and the above

mentioned affecting factors, it is very difficult to realize uniform exit temperaturesamong the individual reforming tubes. A deviation of 30F around targettemperature is quite common.

Normally, operators intend to increase the steam to carbon ratio to avoid hotspots in reformer tubes. Increasing the steam to carbon ratio has three

disadvantages when considering energy consumption:

the steam must be generated in one way or another,

steam must be heated in the reformer using additional fuel gas; and

the steam must be condensed using cooling water before the processgas enters the carbon dioxide removal system.

Process simulations show that increasing the steam to carbon ratio by 0.1 willresult in an increase in energy consumption of 0.15 MMBTU per ton of ammonia.

Secondary Reformer

The effluent from the primary reformer enters the secondary reformer for the final

conversion of methane, generating the required hydrogen for the ammoniasynthesis. Compressed process air is introduced to both combust somemethane and hydrogen in order to supply the required reaction heat for the final

methane conversion and at the same time supply nitrogen gas for the ammoniasynthesis.


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The main objective of the secondary reformer is to adjust air flow to control thedesired hydrogen to nitrogen ratio in the synthesis loop. The hydrogen tonitrogen ratio mainly affects the ammonia conversion in the synthesis loop which

in turn affects the syngas circulation rate, thus the overall energy consumption.Major factors affecting the hydrogen to nitrogen ratio are:


purge gas flow from the synthesis loop,

air flow,

performance of the shift converters,

performance of the carbon dioxide removal section, and

performance of the ammonia converter.

Normally, the ammonia plant simply uses an air to feed gas ratio in the

secondary reformer section to control the hydrogen to nitrogen ratio in thesynthesis loop, but because of the long time lag, the slow effect of the air flow

on the hydrogen to nitrogen ratio and the effects of the downstream sections, itis very difficult to control the ratio at a desired setpoint. The ratio can changefrom 2.5 to 3.5 within two hours, presenting a problem for a smooth and stableoperation. Optimization of the hydrogen to nitrogen ratio to reduce energy

consumption is not an option as long as basic control of the ratio can not bestabilized and maintained.

Shift Converters

The reformed gas effluent consists mainly of hydrogen (H2), carbon monoxide(CO), carbon dioxide (CO2), nitrogen (N2) and steam (H2O). Steam will becondensed in downstream sections. The carbon oxides are poisonous for the

ammonia synthesis catalyst and must be removed before the syngases enter thesynthesis loop. The water gas shift conversion not only converts carbonmonoxide into the more easy removable component CO2, but also generates

additional valuable hydrogen as a by-product. Two converters, the hightemperature shift converter (HTS) and the low temperature shift converter (LTS)are designed to realize this conversion with a guard shift converter (GSC) in

between to protect the low temperature shift catalyst from sulfur poisoning.

The shift reaction is exothermic and while the rate of reaction is favored by

higher temperatures, the final carbon monoxide (CO) content is favored by lowertemperatures. The high temperature shift converter is designed to takeadvantage of the higher reaction rate at elevated temperatures while the low

temperature is designed to take advantage of a lower carbon monoxideequilibrium content at lower temperature.


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Converting the carbon monoxide as much as possible is the main objective ofthis section and this is realized by controlling the inlet gas temperatures to theshift converters.

Major factors affecting the shift converters are:


CO content in the secondary reformer effluent,

HTS inlet temperature,

LTS inlet temperature,

HTS catalyst activity,

LTS catalyst activity, and

steam to gas ratio.

The control objective of this section is to adjust the HTS and LTS inlet

temperature in order to minimize the CO content at the LTS exit. Usually, when

the catalyst is fresh, lower inlet temperatures are feasible, taking advantage ofthe equilibrium position. When the catalyst becomes deactivated, higher inlet

temperatures are required to increase the reaction rate, maintaining the requiredCO conversion.

Due to the deactivation of the catalyst, the operating inlet temperatures of thehigh and low temperature shift converters need to be increased periodically tomaximize the carbon monoxide conversion.

Carbon Dioxide (CO2) Removal

The cooled raw synthesis gas from the shift section, containing about 20%carbon dioxide is fed to the bottom of the CO2 absorber. The carbon dioxide

contained in the gas stream is removed by absorption in an aqueous aminesolution at relatively high pressure and low temperature. The absorbed carbondioxide is subsequently stripped from the amine solution at high temperature and

low pressure in the amine stripper(s).

In this section, the main objective is to minimize the CO2 content in the purifiedsyngas. Amine flow to the absorber and steam flow to the stripper are controlled

to realize this objective.

Major factors affecting the carbon dioxide absorption are:

feed gas flow and composition inlet to the absorber,

lean amine inlet flow to the absorber,

lean amine inlet temperature to the absorber,

lean amine quality to the absorber,


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absorption pressure, and

feed gas inlet temperature to the absorber.

During normal operation, operators often leave the amine flow in manual controlmode and automatically control the stripper overhead temperature by adjusting

the steam flow to the reboiler(s). The carbon dioxide content in the purifiedsyngas varies in the order of 200 PPM around its desired control setpoint andthe energy consumption per ton of carbon dioxide removed is high. Anopportunity exists in this section to optimize the amine flow to the absorber and

steam flow to the stripper(s), reducing energy consumption.

Methanator

A methanator is used to convert all of the remaining carbon oxides (carbon

monoxide and carbon dioxide) through a catalyzed reaction with hydrogen tomethane.

The operation of this section is generally stable. If the methanator experiencesproblems (such as a large temperature increase), it is usually caused by aproblem in an upstream section (such as shift converters or the carbon dioxide

removal section). If such an upset occurs, the upstream sections should bechecked first.

Ammonia Converter

The ammonia converter is one of the most important units in an ammonia plant.

Its performance affects the overall ammonia production and the operating cost

(catalyst, energy consumption, etc.). Usually, one or more ammonia convertersare used depending on the size and the design of the plant. The synthesis gas

distribution has a distinct effect on the overall operation, if more than twoconverters are used.

The objective of the ammonia converter operation is to convert as much of thesynthesis gas as possible to ammonia with minimum hydrogen gas losses (frompurge gas flow).

The temperature profile in an adiabatic or quenched ammonia converter is a

dominant factor in the ammonia conversion. Most ammonia plants operate theammonia converter manually. A temperature variation of 40F is quite common,resulting in an ammonia conversion variation of 0.4%.

Factors affecting the ammonia converter temperature profile are:

synthesis loop pressure,

synthesis gas inlet flow,


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synloop recycle gas flow,

synthesis gas inlet temperature to converter,

synthesis gas inlet composition,

syngas hydrogen to nitrogen ratio,

purge gas flow,

quench flow to each bed, and catalyst activity.

Some of the above factors (such as inert gas composition, hydrogen to nitrogenratio, quench flow to each bed) are adjustable, however they can not be adjusted

instantly. All of the factors are interdependent and a change in one will have aneffect on the others. Consequently, good operation will be a combination of

operating experience and a recognition of the factors affecting the operation ofthe system. A good control system is needed to monitor all the changes andtake proper action whenever a change in process or ambient conditions occurs.

Model Based Advanced Process Control System and Actual Results

Ammonia synthesis is a mature, highly integrated process. Due to the highlyinteractive nature between each process unit, a model based advanced control

system with predictive and adaptive functionality is necessary to optimize theplant operation in achieving the following objectives:

maximum plant profitability considering ammonia, carbon dioxide and

export steam revenues, feed and fuel gas cost, smoother plant operation,

reduced operator intervention by applying closed-loop supervisory

control,

minimum constraint violations with better safety guarantees,

reduced energy consumption,

increased throughput, and

better process performance.

With the model based control system, each process unit has its own closed loopcontrol to respond to the changes in other units. Should a model not function

properly for some reason, such as instrument failure, control valve malfunction,etc, the remaining models still work, continuing to give the plant maximumoptimized operating opportunities.

Based on a typical Kellogg designed, natural gas as feedstock and fuel,ammonia production process, a model based advanced process control system

such as GE Continental Controls, Inc.s MVC

modular multivariable controlsystem can be applied with the following described modules.


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Two modules are set up for the primary reformer section, namely the primaryreformer furnace temperature control module and the primary reformer riser

temperature balance control module. One module is established for thesecondary reformer section, one for the water-gas shift section, one for the CO2

removal section and one module for the ammonia converter. Another modulenamed the integrated module controls the plant pressure profile plus plantthroughputs. All modules are supervised by an overall economic optimizer.

Figure 2 shows the APC block diagram. All control modules and the optimizerwork together to achieve the overall optimization objectives.

Each of the modules controls independently while interactions among theprocess units are taken into consideration.

The model based advanced process control for the ammonia production processis not limited to the above modules. Control modules can be added or deleted

according to the specific ammonia process, such as C.F Brauns, Uhdes, ICIsor Topsoes process. Moreover, the model based advanced control system canbe used in any continuous process

1.

Primary Reformer Furnace Temperature Control Module

This module is set up to control the primary reformer outlet temperature byadjusting the mixed fuel gas pressure to the reformer subject to the maximumtube skin temperature constraint.

The feed gas flow and composition changes, fuel gas heating value (BTU)fluctuations, steam to carbon ratio and mixed gas inlet temperature to the

reformer are considered as disturbance variables.

This module reduces the reformer exit temperatures variations from normal 20 Fto less than 3F, enabling more stable and smoother plant operations. Actualplant operating results from implementing this module in an ammonia plant are

shown in Figure 3.

Primary Reformer Riser Temperature Balance Module

Riser temperatures are a reflection of the reformer tube temperatures. The main

objective of this module is to balance the riser temperatures, thus equalizing thereforming tube temperatures and individual tube methane conversion.

This module will minimize: (1) the temperature differences between successiveriser tube rows, (2) the temperature differences between the first row and the


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last row of the reforming tubes, (3) the temperature difference between the firstrow and the middle row of the reforming tubes. More weight will be placed onthe temperature difference between the tubes in the center and extremes of the

furnace than on the temperature differences between successive tubes. Fuelgas flow controllers are manipulated and optimized to achieve this objective. By

using the predictive and optimizing advanced process control system, thetemperature differences among the reforming tubes have reduced from a normal60F to less than 10F. Plant operating data is shown in Figure 4.

The reduction in riser temperature differences enable plants to raise thereformer exit temperature, increasing ammonia production and to reduce thesteam to carbon ratio. This saves energy while preventing hot spots in the

reforming tubes, thus avoiding carbon deposition to occur. Generally, with thesame reformer exit temperature, the steam to carbon ratio can be lowered by 0.1to 0.2, which results in a 0.15 to 0.3 MMBTU fuel gas savings per ton of

ammonia product.

By avoiding hot spots in reforming tubes, the lifetime of the reforming tubes

and catalyst will increase by two to three years2.

Secondary Reformer Module

In this module, the controller will maintain the hydrogen to nitrogen ratio in theammonia converter inlet at either an operator desired setpoint or a setpointcalculated by the optimizer by manipulating the air flow rate to the secondary

reformer. The air flow is subject to the secondary reformer maximum outlet

temperature constraint. The natural gas feed flow and composition, primaryreformer exit temperature and purge gas flow from the synthesis loop are

considered as disturbance variables.

Because of the predictive and adaptive capability of this module, the hydrogen to

nitrogen ratio is controlled extremely well. The variation of the ratio has reducedtenfold from a normal 0.5 to 0.05. Actual plant data is shown in Figure 4.

Reducing the variation of the hydrogen to nitrogen ratio enables the plant tooperate continuously at the optimum ratio.

It has been reported that the optimum hydrogen to nitrogen ratio is 2.65, but noplant is operated at this low ratio because of synthesis gas compressor, air

compressor and refrigeration system limits. Normally, a ratio of 2.8 to 2.85 willresult in acceptable ammonia conversion and relatively low energy requirementsto maintain circulation. The optimum ratio depends on the plants situation. In

summertime, because of the limitation of the air compressor, the ratio should behigher to achieve maximum throughput while in wintertime the ratio should be


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lowered to yield optimal operation. Stabilizing the hydrogen to nitrogen ratio andoperating at a optimum ratio can save the ammonia plant about 0.2 MMBTU ormore per ton of ammonia product.

Shift Module

This module is designed to minimize the CO content in the reformed gas bycontrolling the outlet temperatures of the high temperature (HTS) and low

temperature shift (LTS) converters by manipulating the inlet temperatures. Feedgas flow and composition are considered as disturbance variables.

Minimizing the CO content from the shift section effluent has severaladvantages:

more hydrogen will be generated by the water gas reaction,

less hydrogen will be used in the downstream methanator section toconvert the carbon monoxide to methane, and

less methane (because of lower CO conversion to methane in the

methanator) will be present in the synthesis loop which results in lesspurge gas losses.

By applying this module, at least two more tons of ammonia per day can beproduced for a typical 1000 TPD ammonia plant.

Carbon Dioxide Removal Module

In this module, the objective is to control the absorber overhead carbon dioxidecomposition by controlling the amine circulation flow rate and steam flow rate to

the stripper bottom reboiler(s), subject to the stripper overhead temperatureconstraint. By regarding the feed gas flow and composition, lean amine inlettemperature to absorber and lean amine quality as disturbance variables, the

variation of CO2 content in purified syngas can be reduced to less than 10 PPM.

Minimizing carbon dioxide content in purified syngas has the following two

advantages:

reduced hydrogen consumption in the methanation section, and

reduced the methane content in make-up syngas.

Ammonia Converter Module

The main objectives of this module are to stabilize and optimize the converter

temperature profile for maximum ammonia conversion by adjusting the quench


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flows to each converter bed. The synthesis loop pressure, inert gascomposition, hydrogen to nitrogen ratio, syngas inlet temperature to converter,syngas feed flow to converter and ammonia content in syngas feed flow are

considered as disturbance variables.

This module can reduce the temperature variations from 20F to less than 4F.The ammonia conversion per pass through the converter can increase about0.2-0.3% which reduces energy consumption by 0.1 to 0.2 MMBTU per ton ofammonia product.

Integrated Module

This module is designed to control the pressure profile and overall throughputsof the plant. Feed gas flow, syngas compressor suction pressure, synthesis loop

pressure, medium steam pressure and steam to carbon ratio are all controlled

variables.

The feed gas flow will be maximized in most cases to maximize ammoniaproduction subject to the following constraints:

feed gas flow maximum controller output,

syngas compressor suction pressure maximum controller output,

air compressor maximum speed,

refrigerant compressor maximum speed, and

primary reformer firebox maximum flue gas pressure.

Medium steam pressure is minimized to yield a higher pressure differential in thehigh pressure stage of the turbine driving the syngas compressor. Thisgenerates more horsepower subject to the maximum controller output constraints

of all medium pressure steam users.

More than ten tons of additional ammonia per day can be produced for a typical

1000 TPD facility by utilizing this module.

Overall Economic Optimizer Module

The real-time, online economic optimizer module maximizes an objectivefunction based on the total ammonia produced, exported carbon dioxide andsteam revenue, fuel gas and feed gas cost.

Economic Optimizer Objective Function:

Profit = NH3 Revenue + CO2 Revenue + Steam Revenue - Feed Cost - Fuel Cost


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The economic optimizer maximizes the plant profit by producing the optimumamount of ammonia product, CO2 product and exported steam. The optimizer

determines the conditions that produce maximum profit subject to constraints asdetailed at the individual module level plus the following constraints:

minimum steam to carbon ratio,

minimum steam export, and

minimum ammonia production.

During each optimizer pass, a new set of manipulated variable setpoints ischosen and are used to calculate the constraints and the objective function. Theoptimizer then passes the setpoints to the corresponding modules. With these

optimized setpoints, the plant is operated at its optimum state, maximum profit isobtained.

Summary

A model based advanced control system enables ammonia plants to operatemore stabile, smoother and more profitable. Due to the advanced control

systems predictive and adaptive capabilities, the interactive effects betweeneach process unit are handled properly and the effect of upset conditions on thedownstream units can be minimized. Furthermore, if one module is not

functioning due to instrument or control valve failure, the remaining modules willwork, giving the plant maximum optimizing operation opportunities.

For a typical 1000 TPD ammonia plant, with the model based advanced processcontrol system, such as GE Continental Controls, Inc.s MVC

multivariable

control system, in excess of twelve tons of additional ammonia can be produced

daily with an energy reduction of 0.6 MMBTU per ton of ammonia produced.Actual operating data is shown in Figure 5 and Figure 6. Depending on thecapacity of the ammonia plant, the payback time of the model based advanced

process control system is four (4) to seven (7) months.


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LITERATURE CITED

1. Lin, R., Viswanathan, S., Poe, W.: Cost Effective Control Solution for

Processing Plants, Fertilizer Focus, September, 1998.2. Twiggs, M.V., Editor, Catalyst Handbook, second edition, Wolfe Publishing,

Ltd., PP 269, 1989.

FUTHER REFERENCES

1. Lin, R., Munsif, H., Poe, W., Primary Reformer Operation: A UniqueApplication of Multivariable Control, Nitrogen No. 230, November-December1997.

2. Grasdal, K., Barone, Peter, Poe, W., Benefits of Advanced Control toAmmonia Plant Operations, AIChE Ammonia Safety Symposium, SanFrancisco, California, September 22-24, 1997.

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