13 replies to this topic

[moein_omg]

Dear Professionals,

 

I have recently been engaged in designing of a propane refrigeration system, the process duty of this sytem is to chill the liquid butane and also propane boil-off gas from storage tank in kettle type heat exchangers. The whole system shall be designed to withstand 23 barg as vapor presure of the propane when the system is shut down and the temperature rises. The compressor used is reciprocating.

now here are my questios:

 

1- Are we forced to consider this design pressure (23 barg) for all the parts of the system. It would mean that all the compressor parts be designed for this pressure which would have high cost effects. Is there any other process solution for it for example SDVs. (I think PSV setting can not be below the 23 barg because lots of the refrigrent will be waisted during every unit shut down and also manual isolation valves are not reliable for determining the design pressure)

 

2- another point, if we raise the design pressure to 23 barg, do we still need to provide all vessels with PSVs sized for fire case. Is it possible to use only one PSV for all the system since there is no shut down valve in the system.

 

by the chance if anyone can provide me with a sample of such system designs, I would be very greatful.

 

many thanks in advance for your kind answers

 

[Art Montemayor]

moein_omg:

 

First, please furnish a detailed PFD (or even better, a P&ID) of what you are describing.  I don’t know why engineers try to imitate poets by using written words to describe a process when a drawing is so much accurate and exact.  It is practically impossible to determine what you mean when you state: “engaged in designing of a propane refrigeration system, the process duty of this sytem is to chill the liquid butane and also propane boil-off gas from storage tank in kettle type heat exchangers”.  Is your system one where you are refrigerating propane OR butane?  Are you doing this by refrigerant evaporation?  Within the storage tank?  Or what?

 

Are we forced to consider this design pressure (23 barg) for all the parts of the system?
How can we know (or guess) that?  You are the one who has received the description of what the “design pressure” involves.  You should ask your boss or your manager what they mean by the “design pressure” and its battery limits.  This may – or may not – involve the extent of the instrumentation and safety devices used.

 

Is it possible to use only one PSV for all the system?
Again, how in heaven’s name can any forum member know the answer if you haven’t shown us HOW the system is arranged and composed (THE PFD OR P&ID, remember?).  Furnish the detailed document requested and our members will all be making recommendations from the same document and description without having to translate what you are trying to state.  Engineering drawings are a universal language.

 

The topic of Propane refrigeration has been one that has been dealt with in many of our past threads in our Forums.  Use our Search Engine and you will find many threads and detailed drawings and calculations on the topic.  Attached is just one of these.  It may not be specifically applied to your application, but it certainly shows you the minimum quality and amount of information required to address and resolve your problem.

 

Await your detailed, engineering response.

 

 Propane Single Stage Refrigeration.xlsx   959.61KB   108 downloads

[moein_omg]

Dear Art Montemayor,
 
attached please find PFD for the aforementioned system.
 
Best Regards

Attached Files

•  PFD.pdf   50.18KB   120 downloads

[Art Montemayor]

moein_omg:

 

Thank you very much for your positive and timely response.  I'm sure our members will now have a clearer understanding of what you are confronting.

 

Replies and comments are sure to be generated.

[S.AHMAD]

Dear Moein,

I personally 100% agree with Mr. Art comments. For  example when you are asking whether 1 PSV is sufficient to cater for the process. We can ONLY positively answer by looking at the P&ID that. 

 

 

If you want good answer, provide more detailed information otrherwise we can only give a very generalized answer.

 

S.Ahmad.

[moein_omg]

thanks a lot for your sincere considerations. I am awaiting your kind responses and comments.

[Art Montemayor]

moein_omg:

 

The following are some preliminary comments based on the PFD you have supplied and your notes and questions:

 

The whole system shall be designed to withstand 23 barg as vapor pressure of the propane when the system is shut down and the temperature rises.
Why the refrigeration system “shall be designed to withstand 23 barg” is not explained.  If this is a mandate, then so shall it be.  I feel just like you: the excessive design pressure will generate additional capital costs.  My experience tells me to design the system to withstand the highest credible vapor pressure it can exert at approximately 50 ^oC.  This is the conservative maximum liquid propane temperature I could expect if it were in Saudi Arabia.  If that is acceptable, then the vapor pressure would be approximately 16 barG.  Refer to the attached workbook, tab “Saturated Propane” to see the corresponding pressure.  If 23 barG is not a mandated figure, then please explain in clear details where this value comes from.  As standard design practice, all refrigeration compressors have to have their 1^st stage cylinders rated for an MAWP equal to at least the maximum design vapor pressure of the refrigerant at the maximum local ambient temperature.

 

Are we forced to consider this design pressure (23 barg) for all the parts of the system?
You are forced to consider the maximum, credible, and achievable pressure to exist at all points within the closed propane refrigerant system that is exposed to liquid propane – initially at any temperature.  This is so because the system – if left alone to equalize its temperature with that of the surrounding environment – will attain a temperature approaching the environment’s temperature and this will generate the related vapor pressure.

 

I think PSV setting cannot be below the 23 barg because lots of the refrigrent (sic.) will be waisted (sic.) during every unit shut down and also manual isolation valves are not reliable for determining the design pressure.
If you use the recommended 16 barG setting, then set the PSVs for a value slightly above that.  There should be NO REFRIGERANT WASTED DURING ANY SHUTDOWN.  I don’t know what you mean by “manual isolation valves are not reliable for determining the design pressure”.  When you shut down the system – whether manually or automatically – the system should be left to equalize its related pressures (unless isolated by block valves, either for maintenance or general purposes).

 

Do we still need to provide all vessels with PSVs sized for fire case?  Is it possible to use only one PSV for all the system since there is no shut down valve in the system?
Without the use of a P&ID, I can categorically (due to experience) tell you that you can’t get by with only one PSV.  You need to employ common sense and practical needs.  You need to use the Refrigerant Accumulator, PK-20701-D-4, as a temporary refrigerant storage vessel whenever you shutdown for maintenance or inspection.  This is the basic design logic used by almost everyone I’ve ever known in the refrigeration business.  From a realistic and practical aspect, you are “locked-in” with the refrigerant you load into the system and you ideally should never divorce yourself from it – or lose it at all (unless you have a major, undesired spill or leakage).  This should be the scope of work that you design the unit under.  If you don’t contaminate, crack, or transform the original propane charge why shouldn’t you be able to use it “forever”?  This is, of course, idealistic.  What really happens in real life is that you have leaks, spills, PSV releases, etc.  but these are only a small percentage of the total charge – not a complete loss or replacement.  During many years of operation, I never lost the charges of Freon, Ammonia, Propane, CO₂.  I added make-up as was needed throughout the years, but never any amount near to even 25% of the original charge.  I know that many other engineers have lost complete charges – but this proved to be incompetence, grave operating and maintenance errors, mistakes, and/or total disorganization.  I have to assume that we all do not tolerate such engineering flaws or mistakes.  You have no excuse for not having shutdown block valves that are essentially 100% leak proof.  If needed, then use double-block-and-bleed types if need be.  But you must have the capability of loading make-up propane “on the run” or while you are on line.  And the moment that you admit to using the propane receiver as storage, then you need an independent PSV there.

 

Please review carefully the attached workbook.  It contains the details of what your PFD fails to show and something that is of utmost importance to know.  That thing is your 2^nd stage compressor KO drum, PK-20701-D-1-B.  This vessel IS NOT A KO DRUM.  It is, in reality, what we called in the refrigeration industry, an "economizer" - a direct-contact heat exchanger that is placed between the 1^st and 2^nd stage to cool the 1^st stage discharge and also serve as an intermediate evaporator that feeds the main refrigerant evaporator.  I give you an example of such a set up using ammonia in the workbook.  note that this vessel requires that the 1^st stage discharge be introduced below the vessel's liquid level for effective direct contact heat transfer.  If the function and the workings of this vessel are not explained in detail, it is very difficult for someone who is not experienced in refrigeration design to understand the PFD.

 

 2-Stage Propane Refrigeration System Rev1.xlsx   1.15MB   66 downloads

[S.AHMAD]

Mr. Art has given good comments  proving his high level of experience in process design. Based on the PFD there is possibility that some of the line and pressure vessel can be rated at lower pressure than 230 psig. However I could not positively identify without the operating conditions and P&ID.

 

Good luck

 

s.ahmad

[Art Montemayor]

All members:

 

I highly apologize for unknowingly uploading a file that was incomplete and failed to convey the subjects discussed in my prior post.  It seems that there is a conflict between the new Excel 10 and all prior versions of Excel and some of the copied sketches and information from prior workbooks did not copy into the target subject workbook.

 

I have now reconstructed the intended workbook and uploaded it into the post where it was intended.  It is now Rev1.

 

Thank you.

[moein_omg]

Dear ART,

 

first of all many thanks for your kind comments and sorry for late reply.

 

now, the process is much more clear to me.

 

1- I came to conclusion that for other parts except pipes and equipments at discharge of compressor, disign pressure of 16 barg is adequate. what about condition that we can not control the pressure at compressor suction drums and heat exchangers (downstream of pressure reduction control valves) and inadvertently the pressure rises. do we need to provide PSV for this senario or increase the design pressure to that of discharge of the compressor (discharge of the reciprocating compressors shall have PSVs).

 

2- another point, if we increase the design pressure of low pressure parts of the system to compressor discharge design pressure, do we still need PSV provision for fire case senarios for each of the pressure vessels and heat exchanges (where they shall all be equipted with inlet/outlet manual block valves for maintenance isolation).

I personally think that we can provide only one PSV for each segment of the system that might automatically be isolated upon process shut down (fail to closed control valves and SDVs) please guide me in this regard.

 

 

Best Regards

[Art Montemayor]

moein_omg:

 

Thank you for your response and I hope that my efforts have been of help to you.

 

I concur with your conclusion of using a design pressure in the order of approximately 16 barg.  However, bear in mind that this is a Design pressure.  Here, what I mean by “design” is a process design pressure – as opposed to a mechanical design pressure.  The difference is important in that any PSV set pressure should comply with the mechanical MAWP and not necessarily with the design pressure (whether mechanical or process).  You are dealing with the objective of protecting personnel and equipment and have to be guided by the MAWP of the equipment that is containing the refrigerant fluid.  We only have your PFD and not a P&ID so I have to assume that what you are trying to describe is a situation where you are shut down and the pressure rises in the suction side of the 1st stage of compression.  Am I correct?  (I hate to keep on making assumptions, but without an as-built P&ID, I have no alternative).

 

If the above is the case, then I also have to assume the pressure rise is due to the refrigerant’s vapor pressure set by the ambient temperature.  I also have no knowledge or indication of where the high pressure may migrate within the system because I have no block or isolation valves indicated.  I assume the suction side pressure could only migrate upstream to a check valve at the high pressure liquid expansion valve and downstream all the way to the high pressure side of the system.  That means that the entire system would be subjected to the pressure increase.  You normally would have the mechanical MAWP of the suction side equipment designed and fabricated to contain this condition (16 barg).  I believe I have already mentioned this in my prior post and you should carefully read and understand what I am referring to or otherwise I will be repeating myself continuously.  The final operating discharge of your compressor cannot be higher than the estimated 16 barg – otherwise you are in a lot of trouble because your condenser is either not working or your cooling medium is well above what it should be.  Coolant failure or other causes should shut down the compressor and system.  Without a detailed P&ID we can only speculate.  Your thorough understanding of the mechanical process and the thermodynamics involved is necessary for you to understand this and your query indicates otherwise.

 

You obviously require a safe means of protecting your personnel and equipment in case of a local fire.  This should only require common sense.  However, besides a PSV you may require a means of automatically draining your liquid refrigerant as well.  You are using a volatile, flammable and combustible fluid and this may be considered as an ultra-hazardous fluid in case of a fire.  This should be scrutinized and heavily discussed and resolved at a Hazop meeting based on the as-built P&ID and local conditions.  Propane is a convenient, efficient, and normally acceptable refrigerant.  HOWEVER, the decision to use this type of refrigerant should have been made taking into consideration some of this refrigerant’s negative characteristics or “tradeoffs”: it is flammable, a combustible, and heavier than air.

 

Your assumption that a PSV alone will be sufficient to save your personnel and equipment may not be correct or wise.  It depends on the scenario, the situation, and the duration of the fire effect.  If you expose equipment steel to a direct fire you are subject to imposing a possible steel failure and ultimate equipment rupture failure.  This is a very serious situation when considering a vessel filled with refrigerant vapor – but even worse when there is a liquid propane refrigerant spill due to equipment failure.

 

Once again, without any detailed P&ID and process description it is difficult to discuss with any accurate and specific comments.  Your local design efforts (depending at what step you find yourself in the project) should have defined - or addressed - a lot of what I have already mentioned and this should have led to specific design details regarding the process equipment specifications and safety protection.  Have I assumed or understood this correctly?  In other words, it is normal engineering practice to process design a refrigeration system and subsequently mechanically design the same system using the temperature and pressure limits indicated or advised in the process design.  The mechanical engineer starts his design where the process engineer terminated his - but he must incorporate the needed temperature and pressure operating ranges indicated by the process engineer.  Therefore, the MAWPs of the mechanical equipment and the vessels involved must already have the necessary values for the process to operate as specified from a process point of view.

[S.AHMAD]

Dear Sir,

Mr. Art has explained the details. I just want to remind you that HAZOPS is essential prior to construction. Normally is conducted once the as-built P&ID available.

 

 

 

 

 

 

[moein_omg]

Dear ART,

 

Many thanks again for your kind replies and I really appreciate you for the time that you devoted to answer my questions.

 

Thanks to your advice, I think now I have a clear mind about the process and safety safeguarding for the BOG system.

 

Regarding the current stage of project, I have to say that we are in the basic stage of the project and the detailed P&IDs and equipment sizes are not  finalized. we even have not freezed the heat and material balances.

 

Regarding the HAZOP studies of the project, as far as I know the meeting shall be held upon finalizing the process documents. Am I right?

 

 

Best Regards

[Art Montemayor]

moein_omg:

 

Good, basic inquiry.  When I managed my projects I always demanded a minimum of three (3) Hazops prior to issuing the P&IDs approved for construction.  Of course, this means that all remedial design questions and issues resulting as a result of the Hazops MUST BE RESOLVED AND APPROVED.  In other words, ALL issues raised in the Hazops must be resolved with the approval of the designers and the ultimate owner-operator before issuing the P&IDs for construction.  This may take some time after the last Hazop is done and that is why I always did my Hazops in stages:

1. The first one at the generation and issue of the preliminary P&IDs;
2. The second one during the development of the piping and general arrangements;
3. The third one at the issuance of the modified P&IDs with comments.

Depending on the size of the project and the steps involved, you may do specific P&IDs over certain high-priority, long delivery items such as compressors, reactors, etc. during the initial phase and decide on whether your project is comfortable with issuing the specifications for quotes to vendors in order to accelerate the ultimate purchase with early delivery.  This means that you must attack such specific items as number of compression stages, types of instruments and capacity controls, ultimate capacities of compressors and heat transfer equipment and its effects on piping, layout, ets.  During the course of doing the total design and procurement, there may be further issues and questions that arise and must be answered.  More Hazops may be required depending if that involves changes to the P&IDs or the required operation.

 

You presently have yet to finalize the Process Design - the basic process heat and material balance is what is the basic supporting specifications for the entire project, so you must do this first and establish a frozen base from which to start the detailed process and equipment design.  Once that is done, the detailed equipment data sheets can be issued for comments and approval.  You may start the P&IDs during this time, but you can't finalize the sizes of the equipment, drivers, heat exchangers, instrumentation, and piping.  But you can do a qualitative, first-order of Hazop on the basic, generic process.

 

It all depends on the size and type of project.